vvEPA
EPRI
TVA
United States
Environmental Protection
Agency
Industrial Environmental Research
Laboratory
Research Triangle Park NC 27711
EPA-600/"'-80-021
February 1980
Electric Power Research
Institute
Fossil Fuel Power
Plants Department
Palo Alto. California 94303
FP-1253
Tennessee Valley Authority
Office of Power
Emission Control
Development Projects
Muscle Shoals, Alabama 35660
ECDP B-6
Preliminary Economic
Analysis of NOX Flue Gas
Treatment Processes
Interagency
Energy/Environment
R&D Program Report
-------
RESEARCH REPORTING SERIES
Research reports of the Office of Research and Development, U.S. Environmental
Protection Agency, have been grouped into nine series. These nine broad cate-
gories were established to facilitate further development and application of en-
vironmental technology. Elimination of traditional grouping was consciously
planned to foster technology transfer and a maximum interface in related fields.
The nine series are:
1. Environmental Health Effects Research
2. Environmental Protection Technology
3. Ecological Research
4. Environmental Monitoring
5. Socioeconomic Environmental Studies
6. Scientific and Technical Assessment Reports (STAR)
7. Interagency Energy-Environment Research and Development
8. "Special" Reports
9. Miscellaneous Reports
This report has been assigned to the INTERAGENCY ENERGY-ENVIRONMENT
RESEARCH AND DEVELOPMENT series. Reports in this series result from the
effort funded under the 17-agency Federal Energy/Environment Research and
Development Program. These studies relate to EPA's mission to protect the public
health and welfare from adverse effects of pollutants associated with energy sys-
tems. The goal of the Program is to assure the rapid development of domestic
energy supplies in an environmentally-compatible manner by providing the nec-
essary environmental data and control technology. Investigations include analy-
ses of the transport of energy-related pollutants and their health and ecological
effects; assessments of, and development of, control technologies for energy
systems; and integrated assessments of a wide range of energy-related environ-
mental issues.
EPA REVIEW NOTICE
This report has been reviewed by the participating Federal Agencies, and approved
for publication. Approval does not signify that the contents necessarily reflect
the views and policies of the Government, nor does mention of trade names or
commercial products constitute endorsement or recommendation for use.
This document is available to the public through the National Technical Informa-
tion Service, Springfield, Virginia 22161.
-------
EPA-600/7-80-021
EPRI FP-1253
TVA ECDP B-6
February 1980
Preliminary Economic Analysis
of NOX Flue Gas
Treatment Processes
by
J.D. Maxwell, T.A. Burnett,
and H.L. Faucett
Tennessee Valley Authority
Office of Power
Emission Control Development Projects
Muscle Shoals, Alabama 35660
EPA Interagency Agreement No. D8-E721-FU
EPA Program Element No. INE624
EPRI Contract No. RP783-3
EPA Project Officer: J.D. Mobley
EPRI Project Monitor: Navin Shah
Industrial Environmental Research Laboratory
Office of Environmental Engineering and Technology
Research Triangle Park, NC 27711
Prepared for
Fossil Fuel Power Plants Department U.S. ENVIRONMENTAL PROTECTION AGENCY
Electric Power Research Institute and Office of Research and Development
Palo Alto, California 94303 Washington, DC 20460
-------
DISCLAIMER
This report was prepared by the Tennessee Valley Authority and has been
reviewed by the Office of Energy, Minerals, and Industry, U.S. Environmental
Protection Agency and the Electric Power Research Institute, and approved
for publication. Approval does not signify that the contents necessarily
reflect the views and policies of the Tennessee Valley Authority, or the
U.S. Environmental Protection Agency, or the Electric Power Research Institute,
nor does mention of trade names or commercial products constitute endorsement
or recommendation for use.
ii
-------
ABSTRACT
A preliminary economic evaluation was made to compare seven flue gas
treatment (FGT) processes for the removal of nitrogen oxides (NOX) from power
plant flue gas. The base case power plant was a new, 500-MW unit burning
3.5% sulfur coal and emitting 600 ppm NOX in the flue gas. The capital
investments and annual revenue requirements for three dry N0x-only removal
processes ranged from $38/kW to $48/kW, and 2.1 mills/kWh to 3.6 mills/kWh
respectively. A dry process for simultaneous removal of both sulfur oxides
(SOX) and NOX had a capital cost of $139/kW and an annual revenue requirement
of 6.6 mills/kWh. Wet SOX-NOX systems, which also removed particulates, had
capital costs and annual revenue requirements ranging from $205/kW to $482/kW
and 12.2 mills/kWh to 19.8 mills/kWh respectively. To permit comparisons
between the NOX control systems, the costs of both a limestone slurry flue
gas desulfurization (FGD) unit and an electrostatic precipitator (ESP) were
added to the cost of the dry N0x-only processes and the cost of an ESP was
added to the cost of the dry SOX-NOX process. The capital investments for
these combined systems based on dry NOX removal (both N0x-only and SOX-NOX
removal) were about $165-$175/kW. The unit revenue requirements for the
combined systems ranged from 7.1 to 8.6 mills/kWh for the dry processes.
Thus it would appear that for the base case power unit, the wet FGT processes
are not economically competitive with either the dry N0x-only or the dry
SOX-NOX systems.
iii
-------
BLANK
iv
-------
CONTENTS
Abstract iii
Figures vii
Tables viii
Abbreviations and Conversion Factors xi
Executive Summary xiii
Introduction 1
Design and Economic Premises . 8
Design Premises 8
Power Plant 8
FGT Process Premises 11
Economic Premises , 19
Capital Investment 19
Revenue Requirements 23
Systems Estimated 28
Classification of NOx Removal Processes .... 28
Asahi Chemical S0x-N0x Process - Wet, Absorption-Reduction .... 32
Process Description 32
Analysis of Processing Subsections 35
IHI SOx-NOx Process - Wet, Oxidation-Absorption-Reduction .... 58
Process Description . 58
Analysis of Processing Subsections 60
Moretana Calcium SOX-NOX Process - Wet, Oxidation-Absorption-
Reduction 77
Process Description .... 77
Analysis of Processing Subsections 79
Hitachi Zosen N0x-0nly Process - Dry, SCR 96
Process Description .... 96
Analysis of Processing Subsections .... 96
Kurabo Knorca N0x-0nly Process - Dry, SCR 101
Process Description 101
Analysis of Processing Subsections . 102
UOP SFGT-N, N0x-0nly Process - Dry, SCR 109
Process Description 1 109
Analysis of Processing Subsections . 110
UOP SFGT-SN, S0x-N0x Process - Dry, SCR 115
Process Description 115
Analysis of Processing Subsections 116
Limestone FGD 127
ESP ' 127
-------
Economic Evaluation and Comparison . 129
Procedure for Estimating the Total Capital Investments ...... 129
Results 130
Asahi S0x-N0x Process 130
IHI S0x-N0x Process ." 132
Moretana Calcium S0x-N0x Process . . 132
Hitachi Zosen N0x-0nly Process 135
Kurabo Knorca NOx-Only Process 135
UOP SFGT-N, NOx-Only Process 138
UOP SFGT-SN, SOx-NOx Process 138
Limestone FGD .... 1 141
ESP 141
Procedure for Estimating the Average Annual Revenue
Requirements 141
Results 142
Asahi SOx-NOx Process 142
IHI SOX-NOX Process 144
Moretana Calcium SOx-NOx Process .... 144
Hitachi Zosen N0x-0nly Process . 144
Kurabo Knorca NOx-Only Process . 148
UOP SFGT-N, NOx-Only Process 148
UOP SFGT-SN, S0x-N0x Process 151
Limestone FGD 153
ESP 153
Overall Comparison 153
Energy Requirements for the Alternative NOx FGT Systems .... 156
Wet Process Gas-Phase Oxidation Versus Absorption of NOX .... 159
Wet SOx-NOx Versus Dry SOx-NOx 160
Wet SOx-NOx Versus SCR-FGD . 160
Dry SOx-NOx Versus SCR-FGD 161
Moving-Bed Reactor Versus Parallel Flow, Fixed-Bed Reactor . . . 162
Conclusions and Recommendations 163
References 166
vi
-------
FIGURES
Number Page
S-l Accuracy range of the estimated capital investment for the
alternative SOx-NOx-PM systems xxiii
S-2 Accuracy range of the estimated annual revenue requirement
for the alternative SOX-NOX-PM systems xxiv
1 Classification system for wet NOx removal processes 31
2 Asahi S0x-N0x process. Flowsheet 36
3 IHI SOX-NOX process. Flowsheet 61
A Moretana Calcium process. Flowsheet . 80
5 Hitachi Zosen process. Flowsheet 97
6 Kurabo Knorca process. Flowsheet 103
7 UOP SFGT-N, NOx-only process. Flowsheet Ill
8 UOP SFGT-SN, SOX-NOX process. Flowsheet 117
vii
-------
TABLES
Number Page
S-l NOX FGT Processes Selected for Evaluation xv
S-2 Capital Investment for the Base Case SOx-NOx-PM
Systems xx
S-3 Annual Revenue Requirements for the Base Case SOX-NOX-PM
Systems xxi
S-4 Accuracy Range of the Estimated Capital Investment and
Annual Revenue Requirement for the SOx-NQx-PM Systems . , . xxii
S-5 Results of Alternative SOX-NOX Comparisons ..«..,.,. xxvi
S-6 Comparison of Energy Requirements for the Various NOx FGT
Processes . xxvii
1 NOx Emissions in the United States , 2
2 NOX Emissions from Selected Stationary Sources in the
United States in 1972 3
3 NOx Emission Standards and Projected Research Objectives
for Large Fossil Fuel-Fired Boilers . 4
4 Current Flue Gas Denitrification Processes . . 6
5 Processes Selected for Further Study in Phase II 7
6 Base Case Coal Composition and Input Flow Rate ....... 9
7 Assumed Power Plant Capacity Schedule 9
8 Flue Gas Composition and Flow Rate at the Economizer
Outlet 10
9 Flue Gas Composition and Flow Rate at Air Heater Outlet ... 10
10 Design Conditions and Removal Efficiencies for the
Prescrubber and the Absorber in the Asahi SOX-NOX
Process 13
11 Design Conditions and Removal Efficiencies for the
Prescrubber and the Absorber in the IHI SC^-NOx Process . . 13
12 Design Conditions and Removal Efficiencies for the
Prescrubber and the Absorber in the Moretana Calcium
SOX-NOX Process 14
13 Design Conditions and Removal Efficiencies for the
Catalytic Reactor in the Hitachi Zosen NOjj-Only Process . . 15
14 Design Conditions and Removal Efficiencies for the
Catalytic Reactor in the Knorca N0x-0nly Process ...... 16
15 Design Conditions and Removal Efficiencies for the
Catalytic Reactor in the UOP SFGT-N, N0x-0nly Process ... 16
16 Design Conditions and Removal Efficiencies for the Reactor
in the UOP SFGT-SN, S0x-N0x Process 17
17 Design Operating Conditions and Removal Efficiencies for the
Presaturator and the Absorber in the Limestone FGD
System 18
vili
-------
TABLES (continued)
Number Page
18 Cost Indices and Projections 20
19 Project Expenditure Schedule . 23
20 Projected 1980 Costs for Raw Materials 24
21 Projected 1980 Unit Cost for Labor 24
22 Projected 1980 Unit Cost for Utilities 25
23 Estimated Annual Maintenance Factor 25
24 Annual Capital Charges for Power Industry Financing .... 26
25 Asahi SOX-NOX Process. Material Balance 37
26 Asahi S0x-N0x Process. Base Case Equipment List
Description and Cost 42
27 IHI SOx-NOx Process. Material Balance 62
28 IHI SOx-NOx Process. Base Case Equipment List
Description and Cost 66
29 Moretana Calcium SOX-NOX Process. Material Balance .... 81
30 Moretana Calcium Process. Base Case Equipment List
Description and Cost 84
31 Hitachi Zosen NOx-Only Process. Material Balance 98
32 Hitachi Zosen N0x-0nly Process. Base Case Equipment List
Description and Cost 99
33 Kurabo Knorca N0x-0nly Process. Material Balance 104
34 Kurabo Knorca NOx-Only Process. Base Case Equipment List
Description and Cost 106
35 UOP SFGT-N,NOx-Only Process. Material Balance 112
36 UOP SFGT-N, N0x-0nly Process. Base Case Equipment List
Description and Cost 113
37 UOP SFGT-SN, S0x-N0x Process. Material Balance ...... 118
38 UOP SFGT-SN, SOX-NOX Process. Base Case Equipment List
Description and Cost 121
39 Asahi S0x-N0x Process. Summary of Estimated Total Capital
Investment 131
40 IHI SOx-NOx Process. Summary of Estimated Total Capital
Investment 133
41 Moretana Calcium Process. Summary of Estimated Total
Capital Investment 134
42 Hitachi Zosen Process. Summary of Estimated Capital
Investment 136
43 Kurabo Knorca Process. Summary of Estimated Capital
Investment 137
44 UOP SFGT-N, N0x-0nly Process. Summary of Estimated Capital
Investment 139
45 UOP SFGT-SN, SOX-NOX Process. Summary of Estimated Capital
Investment .....
46 Asahi SOX-NOX Process. Summary of Annual Revenue Require-
ments - Regulated Utility Economics 143
ix
-------
TABLES (continued)
Number
47 IHI SOX-NOX Process. Summary of Annual Revenue Require-
ments - Regulated Utility Economics 145
48 Moretana Calcium Process. Summary of Annual Revenue
Requirements - Regulated Utility Economics 146
49 Hitachi Zosen Process. Summary of Annual Revenue Require-
ments - Regulated Utility Economics 147
50 Kurabo Knorca Process. Summary of Annual Revenue Require-
ments - Regulated Utility Economics 149
51 UOP SFGT-N, NOx-Only Process. Summary of Annual Revenue
Requirements - Regulated Utility Economics 150
52 UOP SFGT-SN, SOx-NOx Process. Summary of Annual Revenue
Requirements - Regulated Utility Economics 152
53 Total Capital Investment for the Alternative SOx-NOx-PM
Systems 154
54 Total Annual Revenue Requirements for the Alternative SOx-
NOx-PM Systems 155
55 Accuracy Range of the Estimated Capital Investment and
Annual Revenue Requirement for the SOx-NOx-PM Systems . . . 157
56 Comparison of Energy Requirements for the Various NOX FGT
Processes • 158
x
-------
ABBREVIATIONS AND CONVERSION FACTORS
ABBREVIATIONS
abs absolute kW
aft^ actual cubic feet kWh
bbl barrel 1
Btu British thermal unit Ib
°C degrees Celsius m
cm centimeter M
dia diameter rag
ESP electrostatic precipitator min
°F degrees Fahrenheit mm
FGD flue gas desulfurization mol
FGT flue gas treatment MW
ft feet Nm^
ft/sec feet per second ppm
g gram psi
G billion (giga, = 109) SCR
gal gallon sec
gpm gallons per minute SFGT-N
gr grain SFGT-SN
hp horsepower sft^/hr
hr hour
in. inch vol
k thousand wt
kg kilogram yr
kl kiloliter
kilowatt
kilowatthour
liter
pound
meter
million (mega, = 106)
milligram
minute
millimeter
mole
megawatt
normal cubic meter
parts per million
pounds per square inch
selective catalytic reduction
second
Shell-Flue Gas Treatment (NOx-
Shell Flue Gas Treatment (S0x-N0x)
standard cubic feet per hour
standard cubic feet per minute
volume
weight
year
xi
-------
CONVERSION FACTORS
EPA policy is to express all measurements in agency documents in metric
units. Most values in this report are given in English units for the conve-
nience of engineers and other scientists accustomed to using the English
systems. The following conversion factors may be used to provide metric
equivalents.
To convert from
English units
acre
barrel3
British thermal unit
degrees Fahrenheit-32
feet
square feet
cubic feet
feet per minute
cubic feet per minute
gallons
gallons per minute
gallons per thousand actual
cubic feet (50°C)
grains
grains per cubic foot
horsepower
inches
inches of water
pounds
pounds per square inch
standard cubic feet per
minute (60°F)
tons per hour
To metric units
hectare
liters
gram-calories
degrees Celsius
centimeters
square meters
cubic meters
centimeters per
second
cubic meters per
second
liters
liters per second
liters per normal
cubic meter
grams
grams per cubic
meters
kilowatts
centimeters
Pascals (Newton per
square meter)
kilograms
Pascals
normal cubic meters
per hour (0°C)
kilograms per second
Multiply by
0.405
158.97
252
0.5555
30.48
0.0929
0.02832
0.508
0.000472
3.785
0.06308
0.1608
0.0648
2.288
0.746
2.54
249.09
0.4536
6895
1.6077
0.252
a. Forty-two U.S. gallons per barrel.
Other conversion factors
to convert from
dollars
megawatt (electric)
megawatt (electric)
To
yen (Spring 1978)
normal cubic meters
per hour
standard cubic feet
per minute
Multiply by
220
3000
1870
xii
-------
PRELIMINARY ECONOMIC ANALYSIS
OF NO FLUE GAS TREATMENT PROCESSES
x
EXECUTIVE SUMMARY
The existing primary and secondary annual ambient air-quality
standard for nitrogen dioxide (N02) of 100 Vg/m3 is being met in the
United States except for a few locations. However, the 1977 Clean Air
Act Amendments require the promulgation of a short-term NO. ambient air-
quality standard (3 hours or less averaging time) unless it can be
demonstrated that the standard is unnecessary for public health pro-
tection. This short-term ambient standard may require additional con-
trol of nitrogen oxides (NO ) from stationary sources, including electric
power generating facilities? The 1977 Clean Air Act Amendments also
require the installation of the Best Available Control Technology
(BACT) for NO control for new or modified facilities, including electric
generating facilities. Finally, Prevention of Significant Deterioration
(PSD) regulations for N0~ are also required by the 1977 Clean Air Act
Amendments. These requirements of the 1977 Clean Air Act Amendments may
lead to more need for NO emission control for stationary sources,
including electric powerXgenerating facilities in the future.
For the present time, combustion modification techniques under
investigation by the Electric Power Reserrch Institute (EPRI) and the
U.S. Environmental Protection Agency (EPA) may provide sufficient
control of power plant NO emissions. However, the NO control capa-
bilities of these techniques may not be adequate to meet the require-
ments of future regulations and alternative approaches, though poten-
tially less cost-effective, cannot be ruled out. Therefore, NO flue
gas treatment (FGT) processes warrant further investigation.
In a previous report covering the first phase of a multiphase NO
removal FGT process study, a state-of-the-art survey was made of all
known NO FGT processes undergoing development in the United States and
Japan (Technical Assessment of NO Removal Processes for Utility Application,
EPA-600/7-77-127, EPRI AF-568, TV&Y-120).Because low NOemission
limits have been emphasized in Japan in recent years, N0x FGT technology
is more advanced in that country and most of the processes included in
the first phase of the NO process study were developed by Japanese
companies. During the preparation of that report, various Japanese and
American companies were contacted for pertinent technical and economic
information outlining their NO removal processes. This information and
data obtained from other sources were combined to describe the major
xiii
-------
technical and economic aspects of each process. The description for
each process consisted of a process description, status of development,
reported economics, utility and raw material requirements, technical and
environmental considerations, and the. relative advantages and disadvantages.
A major objective of the first phase was to determine candidates for
preliminary economic analysis as a second-phase study.
In this report covering the second phase of the NO process study,
the preliminary economics, comprising total capital investment and
annual revenue requirement estimates, were evaluated for the selected
NO FGT processes. The economics were calculated based on a consistent
ser of design and economic premises that have formed the basis for many
previous flue gas desulfurization (FGD) studies done by the Tennessee
Valley Authority (TVA) . Although various conclusions have been advanced
concerning NO FGT processes, no rigorous economic comparisons have been
made. The primary purpose of this study is to compare the process
economics of the different types of NO control technology on a consis-
tent basis.
Current NO FGT technology involves many types of processes in
different, generally early, stages of development. Selection of proc-
esses for this evaluation was therefore based both on technical and
economic assessment and on the need to evaluate different types of
processes. All the processes evaluated in this study are designed for
90% NO removal but most are in an early stage of development and none
have been extensively demonstrated with coal-fired flue gas.
After the publication of the first-phase report, work was terminated
by the developer of one of the eight processes chosen to undergo prelim-
inary economic analysis and thus the process was omitted from this
report (MON Alkali process). The remaining seven NO FGT processes are
listed in Table S-l according to the specific type or" FGT process they
represent—wet sulfur oxides (SO )-NO , dry SO -NO , or dry NO -only.
Jt Jk Jt A A
The Asahi process uses a sodium sulfite (Na.SO-) solution to absorb
both SO and NO . The NO is directly absorbed witn the aid of a ferrous-
chelating compound. After absorption, the NO is eventually reduced to
molecular nitrogen and vented to the atmosphere. The sulfur dioxide
(SO-) is eventually removed as byproduct gypsum (CaSO,-2HL^') and sodium
sultate (Na2SO,). 4
The process developed by Ishikawajima-Harima Heavy Industries (IHI)
uses ozone injected into the flue gas stream to oxidize nitric oxide
(NO) to N02. The NO. and SO are absorbed from the flue gas into a
calcium-based slurry. The NO and SO are ultimately removed as molecu-
lar nitrogen and byproduct gypsum, respectively.
xiv
-------
TABLE S-l. NO FGT PROCESSES SELECTED FOR EVALUATION
X
Process
Type
Wet S0_.-N0_
x x
Asahi
Ishikawaj ima-Harima
Heavy Industries (IHI)
Moretana Calcium
Absorption-reduction
Oxidation-ab so rption-reduct ion
Oxidation-absorption-reduction
DrySO -NOV
* V A
UOP Shell Flue Gas
Treatment (SFGT-SN)
Dry NO -only
~ " ' X
Hitachi Zosen
Kurabo Knorca
UOP Shell Flue Gas
Treatment (SFGT-N)
Sorption of SO and selective
catalytic reduction of NO
x
Selective catalytic reduction
Selective catalytic reduction
Selective catalytic reduction
In the Moretana Calcium process, NO is oxidized to N09 by injecting
chlorine dioxide (C102) into the flue gas stream. The NO- and SO are
absorbed into a limestone slurry in a Moretana plate tower absorber.
The absorbed N02 and SO form various calcium salts. The SO is ulti-
mately discarded as a m?xed sulfite-sulfate sludge. NO is removed as
ammonium nitrate (NH.NO,,) after reaction of the calciumxnitrates
[Ca(N03>2] with ammonium sul fate [(NH.^SO^].
The Universal Oil Products, Inc., (UOP) Shell Flue Gas Treatment
(SFGT) process for simultaneous SO -NO removal (SFGT-SN) uses a copper
oxide (CuO) acceptor for SO removal. XAs the SO is absorbed, the CuO
is converted to copper sulfate (CuSO,). The sulfate is the catalyst for
the selective catalytic reduction (SCR) of NO to molecular nitrogen and
water by reaction with ammonia (NH») . Subsequent regeneration of the
multiple catalyst beds produces a SOg-rich stream which can be used to
produce various byproducts, e.g., sulfur or sulfuric acid (H_SO,).
This process uses a parallel-flow, fixed-bed reactor located between the
boiler economizer and the boiler air heater.
The three remaining processes control only NO . Each process is
based on SCR, i.e., NH, selectively reacts with NOX in the presence of a
catalyst to form molecular nitrogen and water. The reactors for each of
these processes are located between the boiler economizer and the boiler
air heater. These processes are identified as follows:
xv
-------
1. Hitachi Zosen process: A fixed-bed reactor is used in which the
flue gas passes in parallel flow along the surface of the corrugated-
shaped catalyst cells.
2. Kurabo Knorca process; This process uses a moving-bed reactor.
The flue gas stream is passed in crossflow through the bed of
spherically shaped and slowly moving catalyst. This reactor must
be located downstream from a hot electrostatic precipitator (ESP).
3. UOP SFGT-N process; A parallel passage, fixed-bed reactor is
used in which the catalyst is contained in unit cells and the
flue gas is forced across the face of the catalyst layer.
Since some of these FGT processes, as offered by the various vendors,
remove other potential air pollutants in addition to NO , the following
terminology and equipment additions have been assumed to permit the com-
parison of NO FGT processes on a consistent basis.
X
1. Wet SO -NO ; These processes simultaneously remove both SO and
N0x by absorption into an aqueous solution. Particulate collection
is an inherent part of these processes and is achieved by either
a prescrubber or an ESP and a prescrubber in series. These
processes included all those listed as wet SO -NO in Table S-l.
x x
2- Dry SO -NO ; This process simultaneously removes both SO and
N(P!S0_ removal is achieved by using a dry sorbent while NO
removal Is by SCR. In all economic comparisons an ESP has been
included for particulate matter (PM) control.
3. SCR-FGD; These are the dry, SCR-based, NO -only processes pre-
viously identified in Table S-l. For all economic comparisons an
ESP and a limestone slurry process have been included for PM and
SO- control.
Thus the following comparisons can be made on a consistent set of design
and economic premises.
1. Wet, oxidation-absorption-reduction versus wet, absorption-
reduction .
2. Wet SO -NO versus dry SO -NO .
XX J X X
3. Wet SO -NO versus SCR-FGD.
A X
4. Dry SO -NO versus SCR-FGD.
X. X
5. Moving-bed reactor versus parallel flow, fixed-bed reactor for SCR.
xvi
-------
DESIGN AND ECONOMIC PREMISES
A specific set of design and economic premises was established for
the comparative calculations. The basic design and economic premises
used in this report were established by TVA and EPA for power plant
stack gas emission control studies. Although these premises were
originally designed for FGD studies, appropriate additions and minor
adjustments (primarily dealing with NO control), after joint consulta-
tion between TVA, EPA, and EPRI, were made and the resulting premises
were used for this study.
Design Premises
The power plant assumed as a basis for this study was a new, 500-MW
coal-fired boiler at a midwestern location. The coal had a heating
value of 10,500 Btu/lb and contained 3.5% sulfur and 16% ash. The coal
feed rate of 428,600 Ib/hr was based on a heat rate of 9,000 Btu/kWh.
On-stream time for the boiler was 7,000 hr/yr.
The FGT system included all equipment (based on current technology
and information) necessary to achieve 90% NO removal from flue gas
containing 600 ppm NO and to convert this No to environmentally inert
substances. Other design conditions included a minimum stack gas temper-
ature of 175 F, assumed proven technology (only pumps are spared), and
solids disposal ponds or landfills located one mile from the FGT facility.
When reheat was required an indirect steam reheat system was included.
The wet FGT processes were equipped with a common inlet plenum.
This plenum equally distributed the flue gas from the boiler air heaters
to each of the processing trains. However, the dry FGT processes operate
at higher temperatures (572 F-752 F) and for new boilers, as used in
this study, the FGT system was installed between the economizer and the
boiler air heater. (For retrofit applications the flue gas may have to
be taken from the air heater and reheated before entering the FGT system.)
In addition to preparing the design premises for the NO FGT process
itself, the design premises for an overall FGT system, including PM
removal and FGD, were included to allow for comparisons of the various
dry and wet processes. Because the wet SO -NO processes typically use
a prescrubber for PM removal and achieve greater than 90% SO removal,
the design premises for the FGT system included the following removal
efficiencies.
1. Particulate: 99.5%
2. SO : 90%
x
3. NO : 90%
x
For FGT system comparisons, a high-efficiency ESP and a 90% efficient
limestone "scrubbing" FGD system were included with the NO -only FGT
processes and a high-efficiency ESP was included with the Iry SO -NO
X X
xvii
-------
FGT process. The economics of the limestone FGD system have been pre-
viously calculated based on identical TVA premises. Thus the resulting
capital investments and annual revenue requirements for FGD and ESP were
added to the NO -only FGT processes in-calculations of the economics of
the total package system.
The costs associated with PM disposal are not included, but would
be essentially the same for each process. However, the costs associated
with fly ash handling are included. Also for the dry processes, the
costs which may be required for additional soot blowing or washing of
the air heater are excluded in this estimate. Included in the estimate
of the dry SO -NO process are the incremental cost differences which
are required for the air heater to recover vendor-claimed heat credits.
Economic Premises
Capital investment estimates were based on mid-1979 construction
costs. Current equipment costs were obtained from in-house information
or from equipment vendors. These costs were converted to 1979 dollars
by assuming an escalation rate of 8% annually. The equipment installa-
tion charges were calculated as percentages of the equipment cost deter-
mined by previous TVA experience. Indirect investments such as engineering
design and supervision and architect and engineering expense were calculated
as percentages of the equipment cost. Indirect investments such as
construction expense and contractor fees were calculated as a percentage
of the total direct investment. Royalty fees were included.
Annual revenue requirements were based on mid-1980 operating costs
using average capital charges with a 7,000-hr/yr on-stream time. The
current delivered costs of raw materials (obtained from suppliers in the
Chicago area) were escalated by 5%-10% per year, depending on the expec-
tations of these suppliers. The other annual operating charges were
estimated based on ranges obtained from standard chemical engineering
sources and on previous TVA experience. Byproduct sales credits were
given only for 98% H-SO,, based on the projected 1980 sales price of
this byproduct material. Spent catalyst disposal costs or credits were
applied based on disposal cost or scrap metal value of the base catalyst
support material.
PROCESS DESCRIPTIONS AND SYSTEMS ESTIMATED
As expected, both wet and dry SO -NO processes were more compli-
cated than the NO -only processes. Because the NO -only processes
convert the NO to inert molecular nitrogen by contacting NO with NH~
in the presence of a catalyst, these processes have only two? relatively
small, processing sections: an NH» storage and injection system and a
catalytic reactor. The dry NO -only processes in this study are the UOP
SFGT-N, Kurabo, and Hitachi Zosen processes.
xviii
-------
The UOP SFGT-SN process provides, simultaneous SO -NO control.
This UOP process is unique in that the sorbent for SO* forms a base-
metal sulfate which in combination with NKL reduces NO to molecular
nitrogen. Subsequent regeneration of the multiple catalyst beds using
hydrogen gas from a steam-naphtha reformer yields a concentrated SO^
stream which can be converted to H-SO, in an onsite plant. However,
there are alternatives for hydrogen production, such as coal gasifi-
cation, and for S0_ processing, such as elemental sulfur production with
a Glaus unit.
The wet SO -NO processes absorb both SO and NO in an aqueous
solution to form sulfite-sulfate salts and differing amounts of both
nitrate salts and complex nitrogen-sulfur compounds. The sulfite-
sulfate salts are typically converted into byproduct gypsum and the
complex nitrogen-sulfur compounds are thermally decomposed into molecu-
lar nitrogen and sulfate salts. The nitrate salts are either decomposed
with the complex nitrogen-sulfur compounds or they are concentrated in a
byproduct stream to form a fertilizer material. The wet SO -NO processes
included in this study are the Asahi, IHI, and Moretana CalciumXprocesses.
Process descriptions, detailed flowsheets, material balances, and a
current status of development summary were prepared for each process
based on data supplied by the process vendors. Equipment descriptions
and costs were prepared from the available data for each process.
ECONOMIC EVALUATION AND COMPARISON
The process economics consist of capital investment and annual
revenue requirements for each system. The economic estimates were based
on the material balances and equipment lists, and followed standard
engineering practices. Because of the eclectic nature of the source
data, the simplifying assumptions made, and the necessity of projecting
cost into the future, these preliminary estimates are considered to be
accurate to an overall variation of -20% to +40%.
The coal-fired unit has an output of 500 MW including all system
energy requirements up to and including the ESP and the induced-draft
fans. Thus the 500-MW rating does not include the energy usage of the
FGT or FGD processes. However, throughout this report, for the SO -NO -
PM control systems, the capital investment and annual revenue requirement
are given in dollars per kW and mills per kWh based on a net MW rating.
This net MW rating is calculated by derating the 500-MW output for the
energy usage of the FGT and FGD processes.
Comparison of capital estimates of the FGT processes between any
two reports is very difficult on a dollar per kW or absolute dollar
basis. The FGT processes are primarily sized on flue gas flow rate and
this flow rate may vary significantly even between studies for plants of
similar MW rating. This may be due to several "reasons such as differences
in coal composition, coal heating value, plant heat rate, usage of gross
xix
-------
versus net MW rating, etc. To reduce this difficulty of comparison
between reports, the capital investments of the SO -NO -PM control
systems are also shown on a $/aft /min basis. The flue gas flow rate is
2.146 Maft^/min at the economizer outlet at 705°F.
Capital Investment Estimates
Capital investment is the sum of direct investments, indirect
investments, contingencies, land costs, working capital, and royalty
fees. The total capital investment estimates for the seven NO FGT
processes are shown in Table S-2. Where applicable, the capital invest-
ments for the limestone FGD system and ESP are shown as well as the
resulting capital investment for the combined NO , SO , and PM removal
systems. x x
TABLE S-2. CAPITAL INVESTMENT FOR
THE BASE CASE SO -NO -PM SYSTEMS
x x
Total capital investment
M$
Process
UOP SFGT-N
UOP SFGT-SNa
Kurabo Knorca
Hitachi Zosen
Moretana Calcium
Asahi
IHI
FGT
18.4
67.2
21.2
23.3
88.0
104.9
203.6
FGD
50.4
50.4
50.4
-
-
ESP
10.8
14.6
12.1
10.8
7.2
-
Total
79.6
81.8
83.7
84.5
95.2
104.9
203.6
$/kW
Total
165
169
174
175
205
233
482
$/aftJ/min
Total
37.1
38.1
39.0
39.4
44.4
48.9
94.9
a. Based on hydrogen production from naphtha and H9SO,
from S02. 2 *
byproduct
The capital investments for the various NO -only FGT processes
range from $18M ($38/kW) to $23M ($48/kW) . ForXthe combined SCR-FGD
systems, the capital investments range from $80M ($165/kW) to $85M
($175/kW). For the SO -NO processes the capital investments range from
$67M ($139/kW) to $204H ($?82/kW) .
The capital investment for the SCR-FGD systems range from $165/kW
for the UOP SFGT-N - limestone FGD system to $175/kW for the Hitachi
Zosen - limestone FGD system. Given the uncertainties involved in this
estimate it is impossible to select the best combination based on the
capital investment. The capital investment for the UOP SFGT-SN system
is $169/kW, which is between the $165/kW to $175/kW range for the SCR-
FGD systems. Based on this data, it is impossible to select between the
SCR-FGD and UOP SFGT-SN systems.
xx
-------
Among the wet processes, the IHI process, with a capital investment
of $482/kW, is not economically attractive under the conditions studied.
The Moretana Calcium process, with a capital investment of $205/kW, is
about 24% more costly than the least expensive SCR-FGD system and 21%
more costly than the dry SO -NO system. The capital investment for the
Asahi process is $233/kW, wSichxis about 40% higher than the least
expensive SCR-FGD system and 38% higher than the dry SO -NO system.
Though the capital investments are higher for the Asahi^nd^oretana
Calcium processes, neither can be eliminated on the basis of capital
investment alone.
Annual Revenue Requirement Estimates
The annual revenue requirements consist of direct costs for raw
materials, conversion and services, and indirect costs for capital
charges and overheads. The annual revenue requirements for the seven
NO FGT processes are shown in Table S-3. The total annual revenue
requirements for the limestone FGD process and the ESP as well as the
resulting annual revenue requirements for the combined system are also
shown for those NO FGT processes where an additional FGD process and/or
an ESP is needed to allow a direct comparison. In addition, the unit
revenue requirements in mills/kWh for each system are shown.
TABLE S-3. ANNUAL REVENUE REQUIREMENTS FOR
THE BASE CASE SO -NO -PM SYSTEMS
X X
Annual
revenue requirements,
M$
Equivalent unit
revenue requirements,
mills/kWh
Process
UOP SFGT-N
UOP SFGT-SNa
Kurabo Knorca
Hitachi Zosen
Moretana Calcium
Asahi
IHI
FGT
7.2
22.5
9.3
12.2
38.1
39.8
58.6
FGD
14.7
—
14.7
14.7
—
—
"
ESP
2.2
3.0
2.7
2.2
1.5
-
Total
24.1
25.5
26.7
29.1
39.6
39.8
58.6
Total
7.13
7.53
7.91
8.60 -
12.20
12.63
19.82
a. Based on hydrogen production from naphtha and H-SO, byproduct from
so2.
The annual revenue requirements show a much wider spread in values
for the various processes than was apparent in the total capital invest-
ment estimates. These annual revenue requirements range from $7.2M for
the UOP SFGT-N process to $58.6M for the IHI process. When the costs
xxi
-------
for the limestone FGD process and ESP are added in order to compare
total package systems, the annual revenue requirements increase to
$24.1M for the UOP SFGT-N - limestone FGD system.
Based on the unit revenue requirements, the wet SO -NO processes
do not appear as economically attractive as dry processes for the con-
ditions studied. When compared with the least expensive SCR-FGD system,
the unit revenue requirements for the wet SO -NO systems are about 71%,
77%, and 178% higher for the Moretana Calcium, Asahi, and IHI systems
respectively. These higher unit revenue requirements indicate that the
wet SO -NO systems are not economically competitive with the dry NO
control systems under the conditions used in this evaluation. K
The UOP SFGT-SN process, with the credit for the sale of byproduct
H-SO, at $30/ton, has a unit revenue requirement of 7.53 mills/kWh.
Without this credit the revenue requirements is about 14% higher.
In a comparison of the wet SO -NO processes, the Asahi process,
which is based on the direct absorption of NO , and the Moretana Calcium
process, which uses CIO. as a gas-phase oxidant, have similar unit
revenue requirements. The unit revenue requirement is significantly
higher for the IHI process, which uses ozone as the gas-phase oxidant,
than for the Asahi or Moretana Calcium processes.
Based on an accuracy range of -20% to +40%, the resulting range of
estimated capital investment and annual revenue requirement for the SO -
NO -PM control systems are given in Table S-4. These capital investment
ana annual revenue requirement ranges are also displayed in Figures S-l
and S-2, respectively.
TABLE S-4. ACCURACY RANGE OF THE ESTIMATED CAPITAL INVESTMENT
AND ANNUAL REVENUE REQUIREMENT FOR THE SO -NO -PM SYSTEMS
X 2v
(Based on a -20% to +40% range)
Capital
investment, $/kW
Unit revenue
requirements,
mills/kW
Process
UOP SFGT-N
UOP SFGT-SNa
Kurabo Knorca
Hitachi Zosen
Moretana Calcium
Asahi
IHI
Low
132
135
139
140
164
186
386
Base
165
169
174
175
205
233
482
High
231
237
244
245
287
326
675
Low
5.70
6.02
6.33
6.88
9.76
10.10
15.86
Base
7.13
7.53
7.91
8.60
12.20
12.63
19.82
High
9.98
10.54
11.07
•12.04
17.08
17.68
27.75
a. Based on hydrogen production from naphtha and H2SO, byproduct
from SO
2*
xxii
-------
IHI
t/5
W
u
§
Asahi
Moretana
Calcium
Hitachi
Zosen
Kurabo
Knorca
UOP
SFGT-SN
UOP
SFGT-N
I I I I I I I T
T I
L_L
i
i
i
i
i
i
i
i
120 160 200 240 280
320 360 400 440 480
CAPITAL INVESTMENT, $/kW
520
560 600 640 680
Figure S-l. Accuracy range of the estimated capital investment for
the alternative SO -NO -PM systems (based on a -20% to +40% range).
X X
-------
IHI
Asahi
Moretana
Calcium
en
to
g Hitachi
9 .Zosen
Kurabo
Knorca
UOP
SFGT-SN
UOP
SFGT-N
I I I
LI
i
i
i
12 16 20 24
ANNUAL REVENUE REQUIREMENT, MILLS/kWh
28
32
Figure S-2. Accuracy range of the estimated annual revenue requirement
for the alternative SO -NO -PM systems (based on -20% to +40% range).
X X
-------
The results of these comparisons are summarized in Table S-5. The
wet SO -NO processes do not appear economically attractive for new
power plan? applications when compared with either the dry SO -NO
process or the SCR-FGD systems. Comparisons between the dry UOP SFGT-SN
process and SCR-FGD systems are inconclusive. Comparisons within the
SCR-FGD systems, i.e., moving versus parallel flow, fixed-bed reactors
are also inconclusive. However, recent trends in Japan indicate the
fixed-bed reactor systems are preferred.
Energy Consumption
Because of the likelihood of future increases in the costs of
energy, the energy consumption of FGT methods has become increasingly
important. The energy requirements for each of the seven NO FGT
systems, the ESP's, and the limestone FGD unit were estimated and are
shown in Table S-6. The fuel, steam, electricity, and heat credits are
shown as equivalent MBtu/hr. (The electricity is also shown as equiv-
alent megawatts.) The total energy consumption for each system is shown
as a percent of the boiler capacity.
The dry processes consume the least net energy, 4% of the boiler
capacity. As a result of large heat credits claimed by UOP, the UOP
SFGT-SN process has a slightly lower net energy use, 3.6% of boiler
capacity, than the SCR-FGD systems, where energy use ranges from 3.7% to
3.9% of boiler capacity. These heat credits claimed by UOP consist of
the following: (1) heat of compression from work done on the flue gas
by the flue gas blower, (2) heat of reaction released during acceptance
and regeneration, (3) heat of combustion released to flue gas from
purged regeneration gas remaining in the reactor at the end of regen-
eration, (4) heat of combustion from burning of excess reducing gas, and
(5) extra heat available from the flue gas resulting from removal of
sulfur trioxide (S0_) and subsequent lowering of the dewpoint tempera-
ture. The heat required from the flue gas and regeneration gas for
heating steam and the NH_ air mixture was debited to the UOP SFGT-SN
process. The actual recovery of portions of the heat credits claimed
by UOP to be recovered at the air heater may require additional capital
expenditures. The incremental capital investment for a larger air
heater to recover these heat credits has been included in this report.
The IHI process consumes the most energy (about 19%). The order by
process type in increasing energy costs is dry SOx~NOx, SCR-FGD, and wet
SO -NO . The major energy consumption is electricity for all of the
SOX-NOX processes except the UOP SFGT-SN process, in which fuel is the
largest energy consumption. In most of the processes, equipment drives,
particularly for the flue gas blowers, consume most of the electricity.
However, for the IHI process, most of the electricity is consumed in
generating ozone. Heat credits from the exothermic NO reduction reac-
tions recoverable at the air heater have been given toxthe dry processes
and hence subtracting these from the fuel, steam, and electrical require-
ments give the total energy consumption shown in the table. Although
the UOP SFGT-SN process required incremental air heater costs to recover
xxv
-------
TABLE S-5.
RESULTS OF ALTERNATIVE SO -NO COMPARISONS
x x
Alternatives
The more
attractive alternative
Magnitude of
the difference
Wet absorption-reduction versus wet
oxidation-absorption-reduction (ozone)
Wet absorption-reduction versus wet
oxidation-absorption-reduction (C10-)
' • -~ •«- versus dry SO -NO
Wet SO -NO
Wet SOX-NOX versus SCR-FGD"
Dry SOX-NOX versus SCR-FGD
Moving-bed reactor versus parallel flow,
fixed bed reactor with SCR-FGD
Wet absorption-reduction
Inconclusive
Dry SO -NO
SCR-FGD" x
Inconclusive
Inconclusive
Significant
Insignificant
Significant
Significant
Insignificant
Insignificant
-------
TABLE S-6.
COMPARISON OF ENERGY REQUIREMENTS FOR THE VARIOUS NO FGT PROCESSES'
jj
Process
Limestone FGD
Cold ESP (99.5% efficient)
Cold ESP (99.5% efficient:
UOP SFGT-SN
Cold ESP (94% efficient)
Hot ESP (98% efficient)
Asahi
IHI
Moretana Calcium
Moretana Calcium and ESP
(94%)
Hitachi Zosen
Hitachi Zosen and FGD and ESP
(99.5%)
Kurabo Knorca
Kurabo Knorca and FGD arid ESP
(98%)
UOP SFGT-N
UOP SFGT-N and FGD and ESP
(99.5%)
UOP SFGT-SN
UOP SFGT-SN and ESP
(99.5%)
Fuel, Steam,
MBtu/hr MBtu/hr
78
- -
_ _
- -
- -
216 105
23 232
117
117
1
79
1
79
. 1
79
200 18
200 18
Electricity,
MBtu/hr
74
8
13
4
19
189
581
244
248
20
102
20
113
20
102
91
104
Heat
Total equivalent ,
credit, energy consumption,"
MBtu/hr %
_
-
-
-
-
-
-
-
-
14
14
16
16
15
15
160
160
of boiler capacity
3
0.2
0.3
0.1
0.4
11
19
8
8
0.2
4
0.1
4
0.2
4
3
4
a. Does not include energy requirement represented
b. Based on a 500-MW boiler,
and a boiler efficiency of
a gross heat rate of
by raw materials.
9,000 Btu/kWh
for generation
of electricity,
90% for generation of steam.
-------
heat credits, the SCR-FGD systems do not have the Incremental costs
since the heat credit from just the exothermic NO reduction reaction is
very small. Energy requirements for the purchase! raw material and PM
disposal are not included.
CONCLUSIONS
The following conclusions concerning current SO -NO technology are
subject to numerous important constraints which must oe taken into
consideration in the use of these economic data. These conclusions are
strictly applicable only to the base case, new, 500-MW coal-fired boilers.
Under other power plant applications, particularly retrofit operations,
these conclusions may not be valid. Because of the early status of
development of some of these processes (particularly the wet SO -NO
processes) and the lack of any long-term continuous operational testing
on coal-fired boiler flue gas, the installation and operating costs
could change significantly.
The first and most obvious conclusion is that NO FGT processes
using ozone as a gas-phase oxidant are prohibitively expensive for coal-
fired power units. The total capital investment and the unit revenue
requirements for the IHI process using ozone are more than twice that of
a comparable dry, simultaneous SO -NO system or SCR-FGD system. Ozone
generation accounts for a significant portion of these higher costs (38%
of the capital investment and 10% of the energy consumption is for ozone
generation). Barring some unforeseen major breakthroughs in the manu-
facture of ozone it is highly unlikely that NO FGT systems based on
ozone will be economically attractive in the future.
Although the other wet SO -NO processes (the Asahi process and the
Moretana Calcium process) are more expensive in terms of capital invest-
ment than the dry, simultaneous SO -NO or SCR-FGD systems, the differences
are about the same order of magnitude as the uncertainties surrounding
the cost estimates. Thus, even though the capital investments for the
Asahi process and the Moretana Calcium process are higher, neither can
be eliminated on the basis of the capital investments alone. However,
on the basis of unit revenue requirements, the Asahi and the Moretana
Calcium processes are significantly more expensive than the least expen-
sive dry, SCR-FGD system. Thus, these processes appear economically
unattractive at this time. (It should be noted, however, than in order
to achieve the minimum 90% NO reduction for this study, the Moretana
Calcium and Asahi processes achieve 95% and 99% SO- removal, respectively,
which is greater than the 90% SO,, reduction achieved by the other alterna-
tives. )
A comparison of the dry SO -NO process and the SCR-FGD systems
gives inconclusive results. A more detailed comparison of the two most
promising NO FGT technologies, dry SO -NO and dry NO -only FGT, based
on actual operating experience is needed before the best alternative
xxviii
-------
under various circumstances can be determined. To estimate the cost of
the UOP SFGT-SN process, hydrogen production from naphtha and conversion
of SO- to a H_SO. byproduct are assumed. However, if in the future the
supply of napntha becomes uncertain, it may be necessary to produce
hydrogen by an alternative method, for example, coal gasification.
Also, if no market is available for ^SO, byproduct, SO processing
alternatives may be used, such as elemental sulfur, or liquid SO.
production. Inclusion of alternative processing schemes for the UOP
SFGT-SN process may significantly alter the economic comparisons presented
in this study.
The capital investments of the dry N0x-only FGT processes do not
differ significantly. Variations in catalyst cost account for most of
the differences. The unit revenue requirements of the dry NO -only FGT
systems vary more than the total capital investments, but have some
degree of Overlap within the range of uncertainty associated with pre-
liminary economic studies. Further complicating a comparison of unit
revenue requirements is the fact that the catalyst lifetime, one of the
major annual costs for the dry N0x FGT processes, is uncertain. The
lifetimes estimated by most of the process vendors cannot be verified
because their catalysts have not been used for NO removal on coal-fired
boiler flue gas during a long-term continuous operation. However, with
a catalyst life significantly greater than the 1 year assumed in this
study (the vendors expected 2 to 3 years), the annual revenue require-
ment would be substantially reduced.
The capital investment and revenue requirements for the Kurabo
Knorca system, with a moving-bed reactor, range between those of the UOP
SFGT-N and Hitachi Zosen systems, with parallel flow, fixed-bed reactors.
Therefore no significant distinction can be stated between the moving-
bed and parallel flow, fixed-bed systems. However, recent trends in
Japan indicate the fixed-bed systems are preferred.
The sensitivity of the dry processes to the cost of NH_ has been
the subject of some concern. In this study, the annual NH« cost was
less than 10% of the annual revenue requirement for two out of the three
dry NO FGT processes (Hitachi Zosen and Kurabo Knorca), and the NH_
cost accounted for about 13% of the annual revenue requirement for the
remaining dry process (UOP SFGT-N). Thus, under the premises used in
this study, the annual revenue requirement may be expected to increase
about 10% when the NH» cost is doubled.
The estimates presented in this report are based on currently avail-
able information about the various N0x removal processes. The full impacts
of the NO removal processes on the performance of downstream equipment
(such as wet SO- scrubbers, air heater, ESP or baghouse, and flue gas fan)
are not completely understood at this time. For example, modifications
to air heater design and operation may be necessary to avoid excessive
deposition of ammonium salts. Also, treatment of scrubber byproduct and
blowdown streams may be necessary to prevent release of excessive quantities
of ammonium salts into the environment. Cost data presented in this
xxix
-------
report may have to be revised as the information concerning these impacts
becomes available from future pilot and prototype testing.
Recommendations
Additional preliminary and definitive economic analysis of NO con-
trol systems should be conducted. The results and conclusions of this
study are based on a single base case situation and extrapolation of
these results to other conditions could lead to erroneous conclusions.
Also, current and future development testing will better define the FGT
technology and its impact on downstream equipment, such as the air
heater, FGD system, etc. For these reasons, additional economic evalua-
tions, including both FGT and combustion modifications, should be performed.
Several variations of the base case should be included to determine the
sensitivity of these NO control technologies to design and economic
variations. The case variations should include power unit capacity,
sulfur level in coal, NO removal efficiency, raw material stoichiometry,
and byproduct sales revenue.
XXX
-------
INTRODUCTION
Manmade nitrogen oxide (NOX) emissions are classified as stationary or
mobile, depending upon their source. Each source is responsible for approxi-
mately half of the total U.S. NOX emissions. These sources can be further
divided into the various groups listed in Table 1. The estimated amounts and
the percentages of the total amount of manmade NOX released from each source
for the years 1970, 1972, and 1974 are also shown. During this period the
only two groups with increasing NOx emissions were the mobile sources,
particularly automobiles and trucks, and the fuel combustion group of
stationary sources.
The NOx emission limits for mobile sources were formulated in 1971 by
the U.S. Environmental Protection Agency (EPA), originally for compliance by
the mid-1970's. However, because of technical, economic, and other considera-
tions, compliance with these standards for automobiles has been delayed. Thus
the emphasis on decreasing manmade NOX emissions, at least for the near
future, may be shifted toward the stationary sources. Since the fuel combus-
tion group represents 90-95% of all the emissions from stationary sources,
this group appears likely to undergo more stringent regulation.
The fuel combustion group can be divided into other subgroups according
to the type of fuel used and the type of operation being performed. As one
would expect, the six fuel combustion sources emitting the largest quantities
of NOx are utility and industrial boilers burning coal, oil, or gas. These
six sources and the estimated amount of NOX emitted for each are listed in
Table 2. The major source of NOx in the fuel combustion group is coal-fired
utility boilers, followed by oil-fired industrial boilers and oil-fired
utility boilers; approximately 30% of all stationary-source NOX (about 15%
of the total manmade NOx) is emitted by coal-fired utility boilers.
The projected increasing reliance on coal as the major fuel for the
generation of electrical energy and the gradual phaseout of oil and natural
gas indicate that the NOX contribution from large stationary sources may
increase in the. future. In addition, the anticipated enforcement of regula-
tions requiring the use of coal in new, large industrial boilers and the
conversion of existing large boilers from oil or gas to coal will further
aggravate efforts to control U.S. NOX emissions. In fact, it has been
recently estimated (12) that with a continuation of the present controls on
large utility and industrial boilers, the total annual NOx emissions will
increase about 60% from an estimated 4.7 Mtons in 1975 to 7.6 Mtons in 1985.
With the widespread application of combustion modifications on new stationary
sources, the total amount of NOX emitted by large boilers may increase about
40% to nearly 6.6 Mtons in 1985 (12).
-------
TABLE 1. NOX EMISSIONS IN THE UNITED STATES
Annual NOX emissions
1970
Source
Mobile
Highway
Nonhighway
Subtotal
Stationary
Fuel combustion
Industrial processes
(non combustion)
Solid waste
Miscellaneous
Subtotal
Total NOX emissions
Amount,
Mtons
6.9
2.4
9.3
10.1
0.6
0.3
0.1
11.1
20.4
% of total
NOX
33.8
11.3
45.6
49.5
2.9
1.5
0.5
54.4
100.0
1972
Amount ,
Mtons
7.9
2.6
10.5
10.8
0.6
0.2
0.1
11.7
22.2
% of total
NOX
35.6
11.7
47.3
48.6
2.7
0.9
0.5
52.7
100.0
1974
Amount ,
Mtons
8.1
2.6
10.7
11.0
0.6
0.1
0.1
11.8
22.5
% of total
NOX
36.0
11.6
47.6
48.9
2.7
0.4
0.4
52.4
100.0
Source: (19)
-------
TABLE 2. NOX EMISSIONS FROM SELECTED STATIONARY SOURCES
IN THE UNITED STATES IN 1972
NOx emissions
Annual amount, % of stationary % of manmade
Stationary source short tons sources NOx
Utility boilers
Coal-fired 3,853,000 30.7 15.6
Oil-fired 1,228,000 9.8 5.0
Gas-fired 920,000 7.3 3.7
Industrial boilers
Coal-fired
Oil-fired
Gas- fired
Total
810,000
1,372,000
541,000
8,724,000
6.5
11.0
4.3
69.6
3.3
5.6
2.2
35.4
Source: (15)
Following enactment of the Clean Air Act of 1970, EPA released New
Source Performance Standards (NSPS) for NOX emissions in December 1971. The
latest NSPS were released by EPA in June 1979. Both of these sets of stan-
dards for large boilers (>250 MBtu/hr, 28-MW equiv) are shown in Table 3.
The most recent standards are based on nitrogen in fuels (i.e., ranging from
0.2 Ib NOx/MBtu for natural gas with no bound nitrogen to 0.6 Ib NOx/MBtu
for coal which contains about 1% bound nitrogen), the size of the utility
boiler, and the best available technology for NOX control. However, the
NSPS may become more stringent as NOX control technology is further developed
(see projected research objectives in Tablt 3). The need for additional con-
trols may also increase as the amount of NOX emitted to the atmosphere
increases.
NOX EMISSION CONTROL
Although NOX is present in relatively minor amounts in the flue gas, a
single, 500-MW coal-fired power plant releases approximately 10,800 tons/yr
of NOX to the atmosphere. The NOx in the flue gas is generated during com-
bustion from the oxidation of nitrogen-containing compounds in the fuel
(fuel NOX) and the thermal reaction of molecular nitrogen and oxygen during
combustion (thermal NOX). The portion of the total NOX emissions from each
of these sources is a function of many variables and is quite site specific.
Of the two types of NOX control technology currently under development,
combustion modifications and flue gas treatment (FGT), combustion modifica-
tions will probably receive the initial emphasis since it is considered the
most cost-effective method of controlling NOX emissions (12).
-------
TABLE 3. NOX EMISSION STANDARDS AND PROJECTED RESEARCH OBJECTIVES
FOR LARGE FOSSIL FUEL-FIRED BOILERSa
1971 EPA
standard (18)
Gaseous fuel
Liquid fuel
Solid fuel
Lb NOx/MBtu
input
to boilerb
0.2
0.3
0.7
NOx
ppmc
150
225
550
June 197.9 NSPS .(20)
Lb NOx/MBtu
input
to boiler
0.2
0.3
0.5 (subbituminous)
0.6 (bituminous)
Projected
research
objectives (15)
1985 .
NOX
ppm°
50
90
100
a. >250 MBtu/hr heat input.
b. Expressed as N02-
c. Calculated at 3% excess ©2, dry basis.
Combustion modifications have their greatest effect on oil- and gas-
fired boilers, where most of the NOx is generated by the thermal mechanism,
and less of an effect on coal-fired boilers, where most of the NOx is formed
from fuel nitrogen. With the current emphasis on coal as the fuel for all
new, large boilers, combustion modifications may not he adequate in control-
ling the amount of NOx emitted by stationary sources. To achieve the reduc-
tion of NOx emissions which may become necessary in the future, FGT may be
required.
PROJECT SCOPE
The primary purposes of this study are to compare the economics of
selected types of NOx FGT processes using a set of consistent premises and
to provide current and more detailed information on the processes. This
report consists of: (1) the design and economic premises, (2) a description
of the processes and process combinations used in the economic evaluations,
and (3) results of the economic evaluations and economic comparisons of
systems.
Process Selection
The initial Phase I NOX report (Technical Assessment of NOx Removal
Processes for Utility Application. EPA-600/7-77-127, EPRI AF-568, TVA Y-120),
completed in late 1977, was a technical assessment of some 48 FGT processes
currently listed in the literature. Each process was defined and evaluated
by the following criteria: (1) process description, including a block flow
diagram; (2) status of development; (3) background of the process developer;
-------
(4) published economic data; (5) raw material, energy, and operation require-
ments; (6) technical considerations; (7) environmental considerations;
(8) critical data gaps and poorly understood phenomena; and (9) overall
technical advantages and disadvantages. In addition, each of the NOX FGT
processes was categorized as a wet- or a dry-treatment process and as a
particular type of wet or dry process. These 48 processes described in the
Phase I study are listed alphabetically in Table 4 as a wet or dry process
and further as a particular wet or dry type. As seen from this table, a
large proportion of the current FGT processes undergoing development belong
to the dry selective catalytic reduction (SCR) type.
One purpose of the Phase I study was to select eight of the most
promising FGT processes for preliminary economic analysis in this phase of
the NOX process study. These processes were selected on the basis of techni-
cal and economic considerations, development status, and the need for a
representative sample.
The processes selected by these criteria are shown in Table 5. Four
dry SCR processes were selected primarily on the basis of type of reactor
(which is the only major piece of process equipment in dry N0x-only systems)
and because this type represents almost half of the processes under develop-
ment. The Universal Oil Products, Inc., (UOP) Shell Flue Gas Treatment
process has two variations: a simultaneous sulfur-oxide - nitrogen-oxide
(SOX-NOX) process (SFGT-SN) and an NOx-only process (SFGT-N), both of which
were included. The two wet oxidation-absorption-reduction processes use
different oxidants with potentially large cost differences. The Asahi
Chemical process was selected as representative of wet absorption-reduction
systems. The MON Alkali process, the eighth process recommended in the
Phase I study, was deleted from this study because development work on the
MON Alkali process was abandoned after publication of the Phase I report.
The JGC Paranox process was considered for substitution as an eighth process.
This process was investigated and correspondence ensued with UOP, the U.S.
licensee. However, since this process technology is almost identical with
the UOP SFGT-N process (the catalyst composition being the only difference),
and the costs are claimed to be about the same as the UOP SFGT-N process, the
JGC Paranox process was also omitted from this report.
To permit comparisons of N0x-only and SOX-NOX FGT systems and to provide
comparisons under a wider range of possible applications, a limestone scrubber
flue gas desulfurization (FGD) unit is also included in the comparisons.
This is used in conjunction with the NOx-only FGT systems in comparisons with
the SOx-NOx FGT systems. Also, the type and cost of particulate collection
required will vary according to the type of FGT system. Therefore, to
provide an equitable comparison between the NOx FGT alternatives, particulate
collection is also included in the comparisons.
Process Descriptions and Evaluations
The process dercriptions in this report are amplified and updated
versions of the material in the Phase I report (EPA-600/7-77-127; EPRI
-------
TABLE 4. CURRENT FLUE GAS DEVITRIFICATION PROCESSES
Dry processes
Wet processes
Selective catalytic reduction
Asahi Glassa
Eneron
Exxon
Hitachi, Ltd.
Hitachi Zosen
JGC Paranox
Kobe Steel
Kurabo Knorca
Kureha
Mitsubishi Heavy Industries
Mitsubishi Kakoki Kaisha
Mitsubishi Petrochemical
Mitsui Engineering and
Shipbuilding
Mitsui Toatsu
Nippon Kokan3
Sumitomo Chemical
Sumitomo Heavy Industries^
Takeda
Ube
Universal Oil Products-Shell1*
Nonselective catalytic reduction
The Ralph M. Parsons
Selective noncatalytic reduction
Exxon Thermal Denox
Adsorption
Foster Wheeler-Bergbau Forschung
Radiation
Ebara-JAERI
Uncertain
Princeton Chemical Research3
Absorption-reduction
Asahi Chemical
Chisso Engineering
Kureha
Mitsui Engineering and
Shipbuilding
Pittsburgh Environmental and
Energy Systems
Oxidation-absorption-reduction
Chiyoda Thoroughbred 102
Ishikawajima-Harima Heavy
Industries
Mitsubishi Heavy Industries
Moretana Calcium
Moretana Sodium
Osaka Sodaa
Shirogene3
Absorption-oxidation
Hodogayaa
Kobe Steel
MON Alkali Permanganate
Nissan Engineering
Oxidation-absorption
Kawasaki Heavy Industries
Tokyo Electric-Mitsubishi HI
Ube
a. Essentially no information available.
b. These companies have two dry processes:
(2)
(1) N0x-only and
-------
TABLE 5. PROCESSES SELECTED FOR FURTHER STUDY IN PHASE II
Process
Type of_proces_s (classification)^
UOP Shell Flue Gas Treatment
(SFGT-N)
Hitachi Zosen
Kurabo Knorca
UOP Shell Flue Gas Treatment
(SFGT-SN)
Moretana Calcium
Ishikawajima-Harima Heavy
Industries (IHI)
Asahi Chemical
Dry NOX only (SCR)
Dry NOX only (SCR)
Dry NOX only (SCR)
Dry simultaneous S0x-N0x (SCR)
Wet simultaneous S0x-N0x
(oxidation-absorption-reduction)
Wet simultaneous S0x-N0x
(oKidation-absorption-reduction)
Wet simultaneous S0x-N0x
(absorption-reduction)
AF-568; TVA Y-120). In addition, flow diagrams and material balances are
included. The design and economic premises are similar to those used in
previous Tennessee Valley Authority (TVA) FGD studies, with the inclusion
of additional premises for NOX FGT. These premises define design and
operating conditions for a representative power unit, permitting comparison
of the systems evaluated on an equal basis.
The economic evaluation consists of comparisons of estimated total
capital investment and average annual revenue requirements for each process
and process combination. Estimates were prepared and comparisons made for
the following systems: (1) wet versus dry S0x-N0x FGT systems, (2) dry NOX-
only systems with limestone FGD (SCR-FGD) versus wet S0x-N0x FGT systems,
(3) SCR-FGD versus dry SOX-NOX FGT systems, (4) wet SOx~NOx FGT systems
using oxidation-absorption-reduction steps versus those using absorption-
reduction steps, and (5) moving-bed reactor versus parallel flow, fixed-
bed reactors for dry NOX FGT.
-------
DESIGN AND ECONOMIC PREMISES
An accurate comparison of the various types of NOx FGT processes can be
made only if a specific set of design and economic premises is established.as
a basis for the design and economic calculations. In addition, an analysis
of various NOx FGT processes is complicated because the dry SCR processes are
NOx-only FGT systems. Therefore, the NOx-only systems have been combined
with a conventional limestone FGD system so that all the alternatives are
compared on an equal basis (i.e., minimum 90% removal of both SQx and NOx).
In addition, particulate collection, the type and cost of which vary between
the different NOx FGT processes, is included for an equitable, overall
comparison of SOx-NOx-particulate matter (PM) control systems. The pertinent
design and economic premises concerning the operation of the power plant,
the FGD system, the various electrostatic precipitators (ESP), and the
FGT systems are discussed in this section.
DESIGN PREMISES
The basic design premises for power plant stack gas emission control
studies were established by TVA and EPA for work conducted prior to this
study. These premises are similar to those used in other technical and
economic evaluations of power plant emission- control technology. After
consultation between TVA, EPA, and Electric Power Research Institute (EPRI),
these premises, with appropriate additions and minor adjustments (primarily
dealing with NOx control), were also used for this evaluation. Using these
general design premises, a set of base case design premises was established.
Power Plant
Plant Size and Fuel—
The only power plant evaluated in this study was a new, 500-MW coal-fired
boiler. A midwestern location was assumed as a basis for the power unit
because of the concentration of power stations in that area and the proximity
of this area to the major coal fields. The fuel for the plant was a coal
having a heating value of 10,500 Btu/lb and containing 3.5% sulfur and 16%
ash. The coal composition and the input coal requirements (based on a heat
rate of 9,000 Btu/kWh) for the 500-MW boiler are listed in Table 6.
-------
TABLE 6. BASE CASE COAL COMPOSITION
AND INPUT FLOW RATE
(500-MW new unit, 9,000 Btu/kWh heat rate)
Wt %» as fired Lb/hr
c
H2
N£
02
S
Cl
Ash
H20
57.56
4.14
1.29
7.00
3.12
0.15
16.00
10.74
246,800
17,700
5,500
30,000
13,400
600
68,600
46,000
Total 100.00 428,600
Power Unit Life, Operating Time, and Capacity-
Based on guidelines suggested by the then Federal Power Commission (FPC)
(7), a new, coal-fired power plant is estimated to have a 30-year life.
However, the operating capacity for these plants varies widely over their
lifetime. Power unit operating schedules that reflect TVA experience were
used in this study. These are shown in Table 7.
TABLE 7. ASSUMED POWER PLANT CAPACITY SCHEDULE
Operating Capacity factor, % Annual operating
yr (nameplate rating) time, hr
1-10 80 7,000
11-15 57 5,000
16-20 40 3,500
21-30 17 1,500
Average for
30-yr life 48.5 4,250
Flue Gas Composition—
Because dry SCR processes operate at higher temperatures (570-780°F)
than wet processes, most process developers recommend that the flue gas be
removed at the outlet of the boiler economizer for passage through the FGT
system. Flue gas for the wet FGT processes is removed at the outlet of the
boiler air heater. Because of air leakage in the boiler air heater—assumed
-------
to be 13% of the stoichiometric combustion air requirements—the flue gas
composition at this position is different than the composition of the flue
gas at the economizer outlet. The flue gas composition at the economizer
outlet is shown in Table 8 and at the boiler air heater outlet in Table 9.
TABLE 8. FLUE GAS COMPOSITION AND FLOW
RATE AT THE ECONOMIZER OUTLET
(500-MW new unit, coal-fired, 9,000 Btu/kWh,
3.5% S, 10,500 Btu/lb HHV as fired,
2,146,000 aft3/min at 705°F)
Component
N2
02
C02
S02
S03
NOX
HC1
H20
Vol. %
73.39
3.23
13.56
0.26
0.003
0.064
0.012
9.48
Lb/hr
3,113,000
156,500
904,200
25,130
317
3,009
661
258,600
Total 100.00 4,461,000 (approx)
TABLE 9. FLUE GAS COMPOSITION AND FLOW
RATE AT AIR HEATER OUTLET
(500-MW new unit, coal-fired, 9,000 Btu/kWh,
3.5% S, 10,500 Btu/lb HHV as fired,
1,543,000 aft3/min at 300°F)
Component
N2
02
CO 2
S02
S03
NOx
HC1
H20
Vol, %
73.76
4.83
12.31
0.24
0.0024
0.060
0.01
8.79
Lb/hr
3,450,000
258,200
904,200
25,130
317
3,009
661
264,500
Total 100.00
4,906,000 (approx)
10
-------
The flue gas compositions are calculated on the "basis that 95% of the
sulfur in the coal is emitted in the flue gas as SOX [99% sulfur dioxide (S02)
and 1% sulfur trioxide (803)] and 80% of the ash is emitted in the flue gas as
particulates. These values represent TVA operating experience with frontal-
fired coal-burning units.
The NOx concentration of the flue gas at both positions is 600 ppm by
volume and consists of about 95% nitric oxide (NO) and 5% nitrogen dioxide
(N02>. This concentration and composition is believed representative of the
range of concentrations likely to occur in particular applications.
Emission Standards—
Although the pre-1979 NSPS for new, coal-fired utility boilers were in
Ib/MBtu heat input the 1977 amendments to the Clean Air Act of 1970 required
establishment of NSPS that specify a percent reduction as well as a maximum
emission rate with respect to S02, which was accomplished in the 1979 NSPS.
Therefore, each of the FGT process packages in this study includes all the
equipment required to achieve particulate, SOx, and NOX removal efficiencies
of 99.5%, 90%, and 90% respectively. Since the wet FGT processes are capable
of removing more than 90% of the inlet SOX, a limestone FGD system capable
of 90% SOX removal was coupled with the NOx-only FGT systems to form
comparable overall systems. An ESP system was coupled with certain FGT
systems to form a comparable overall system with the wet FGT processes using
wet scrubbers for particulate and chloride removal.
FGT Process Premises
For this study the ducts from the boiler air heaters are assumed to
exhaust to a common plenum connecting the scrubbing trains. Separate ducts
from the plenum to each scrubbing train are equipped with dampers for
individual scrubber shutoff during periods of maintenance or power plant
turndown.
With the exception of the Asahi process where flue gas from an oil-fired
heater is available, and the UOP SFGT-SN process which is a dry S0x-N0x
system upstream of the air heater-stack gas reheat to 175°F for plume
buoyancy is provided by indirect steam reheat of the cleaned gas. For reheat
calculations the amount of entrained water in the flue gas to the reheater
Is assumed to be 0.1 weight percent of the wet-gas mass flow rate in each
processing system.
Separate induced draft (ID) fans are not included for the dry FGT
systems. Rather, a larger boiler ID fan is used to compensate for the
higher pressure drop of the combined boiler-FGT system and the increased
costs are assigned to the FGT process. Separate booster fans and ductwork
for the SOX-NOX systems and the limestone FGD unit are included because of
the larger pressure drops involved with these systems.
11
-------
For systems requiring disposal of sludge or solid waste other than fly
ash, a disposal pond or landfill located 1 mile from the plant site is pro-
vided. These are designed to minimize land and construction costs and are
sized for a 30-year life at the waste flow rates of the process.
In addition, the design premises assume proven technology, not first-of-
a-kind installatxons. Other than spares for pumps no redundancy features
are included. In actual operation, additional spares or redundancy may be
required.
There are certain items which are excluded in this report. The invest-
ment and revenue requirement for the fly ash pond, which should be similar
for each process, are omitted. Also excluded are the costs associated with
additional soot blowing which may be necessary for the air heater with the
dry FGT processes.
Costs for the ash-handling system to move the fly ash to final disposal
and soot blowing of the dry FGT process reactors are included. Also included
are any incremental cost differences required to recover claimed heat credits
by the dry processes, for example, larger air heaters.
Two reactor trains were selected for the dry processes, except for the
Kurabo Knorca system in which four trains were recommended, to provide a
comparative basis.
Asahi Process—
The Asahi process is an example of the wet, simultaneous SOX-NOX>
absorption-reduction type FGT system. The relatively low mass transfer co-
efficients associated with the absorption of NO by aqueous solutions make
large absorbers necessary. Thus the initial -sections of the Asahi process,
prescrubbing and absorption, consist of four separate trains for a 500-MW
coal-fired boiler. The prescrubbers are modified venturi scrubbers and the
absorbers are packed towers. The design conditions and removal efficiencies
in the prescrubbers and absorbers are shown in Table 10.
Although there are four trains of prescrubbers and absorbers, most of
the regeneration section is a single train. To avoid potential forced boiler
outages due to the regeneration section going off-line for routine maintenance
or repairs, excess surge capacity (usually 8 hours) is included at strategic
locations throughout the regeneration system. The process design includes
all equipment needed to dispose of the byproduct gypsum in a landfill. Sodium
sulfate (^2804) is assumed to be sold at no profit.
Ishikawaj ima-Haritna Heavy Industries (IHI) Process—
The IHI process is a wet, simultaneous SOX-NOX, oxidation-absorption-
reduction type FGT system. In this type of FGT process, a gas-phase oxidant
is used to convert NO to N02« Since N02 is absorbed much more readily than
NO, the developers of the IHI process believe that only a single train of
processing equipment for a 500-rMW coal-fired boiler is required. However, to
prevent boiler outages due to shutdown of the FGT system, additional surge
volume (usually 8 hours) has been included in each section of the regenera-
tion system.
12
-------
TABLE 10. DESIGN CONDITIONS AND REMOVAL EFFICIENCIES FOR THE
PRESCRUBBER AND THE ABSORBER IN THE ASAHI SOX-NOX PROCESS
Item
Prescrubber
Absorber
Type
Operating conditions
Pressure drop, in. of H20
L/G, gal/kaft3
PH
Removal efficiency, %a
S02
803
NOX
HC1
Ash
Modified venturi Tellerette packed-bed
7.9
6.2
1.0
10.0
25.0
99.0
99.0
7.9
62.2
6.0-6.5
99.0
86.0
90.0
90.0
90.0
a. [(Inlet-outlet)/inlet]100.
The single prescrubber and absorber have the design conditions and
removal efficiencies shown in Table 11. Essentially all of the absorbed SOx
is converted to gypsum. From initial development work it appears that
approximately 80% of the absorbed NOX is converted to complex nitrogen-sulfur
compounds, 15% is converted to molecular nitrogen, and the remaining 5%
remains as nitrite-nitrate salts. Therefore these conversion rates were
assumed for this study.
TABLE 11. DESIGN CONDITIONS AND REMOVAL EFFICIENCIES FOR
THE PRESCRUBBER AND THE ABSORBER IN THE IHI SOx-NOx PROCESS
Prescrubber
Absorber
.Type
Operating conditions
Pressure drop, in. of H20
L/G, gal/kaft3 '
pH
Removal efficiency, %b
Modified venturi
12.0
NAa
1.0-2.0
803
Ash
95.0
99.5
Spray tower
4.0
NA*
4.0-5.0
90.0
90.0
0.0
0.0
a. Not available, proprietary information.
b. [(Inlet-outlet)/inlet]100.
13
-------
The ozone for the IHI process is generated by corona discharge which
produces 1.0 weight percent ozone in the airstream from the generators. The
ozone generation system capacity is based on a required ozone to NOx mol
ratio of 0.85:1.0 (assuming 5% of the NOX is already N02) in the flue gas.
The process design includes a landfill site for the disposal of the byproduct
gypsum and mixed salts.
Moretana Calcium Process—
The Moretana Calcium process is also a wet, simultaneous SOx~NOx,
oxidation-absorption-reduction process. Although the developers of the IHI
process believe only one train is needed for a 500-MW coal-fired boiler,
Sumitomo Metal, developer of the Moretana Calcium process, recommends two
trains of equipment. Even though the process developers recommend two trains,
only one train of limestone preparation equipment is used and sufficient
surge storage is designed into the limestone slurry tank.
The Moretana tower absorber is designed for the operating conditions and
removal efficiencies given in Table 12. Most of the SOX absorbed is converted
to a mixed sulfite-sulfate sludge similar to that produced in a conventional
limestone FGD system. The NOX, on the other hand, is oxidized to both N02
and nitric acid (HNC-3) and after absorption is converted to molecular nitrogen
and calcium nitrate [Ca(N03>2] respectively. This conversion of NOx to nitrogen
and Ca(N03>2 is assumed to be equimolar, with no other nitrogen byproducts
formed.
TABLE 12. DESIGN CONDITIONS AND REMOVAL EFFICIENCIES FOR THE PRESCRUBBER
AND THE ABSORBER IN THE MORETANA CALCIUM SOx-NOx PROCESS
Item Prescrubber Absorber
Type
Operating conditions
Pressure drop, in. of
L/G, gal/kaft3
PH
Removal efficiencies, %a
S02
SOs
NOX
HC1
Ashb
C102
C12
HC1
HN03
Moretana tower
H20 3.9
30.8
1.0
-
50.0
-
95.0
90.0
-
-
-
—
Moretana tower
8.3
87.1
5.0
95.0
50.0
90.0
100.0
80.0
100.0
100.0
100.0
100.0
a. [(Inlet-outlet) /inlet] 100.
b. Moretana Calcium process is equipped with a 94%
prescrubber.
efficient ESP prior to
-------
The design mol ratio of chlorine dioxide (C102) to NOX is 0.55:1.0 for
90% NOX removal. Equipment is included in each C102 generating unit for
adequate surge storage of C102- The process design also includes a sludge
pond for the disposal of sulfite-sulfate sludge.
Hitachi Zosen Process—
The Hitachi Zosen process is a dry, NOx-only, SCR system with two
processing trains and using a fixed-bed reactor and corrugated-shaped
catalyst. The ammonia (NHs) is diluted with air to 5% by volume for easier
flow control, enhanced mixing, and because the limits of inflammability for
NH3 in air are 15.5 to 27.0%. The design conditions and removal efficiency
in the reactor are shown in Table 13. The overall process design for the
Hitachi Zosen system includes a limestone FGD system and a high-efficiency
cold ESP.
TABLE 13. DESIGN CONDITIONS AND REMOVAL EFFICIENCIES FOR THE
CATALYTIC REACTOR IN THE HITACHI ZOSEN NOX-ONLY PROCESS
_ Item _ Reactor _
Type Fixed -bed (corrugated)
Operating conditions
Pressure drop, in* of H20 2.0-3.0
Space velocity, ft3/hr/ft2 5,000-10,000
NH3:NOx mol ratio (1.0-1.2) :1.0
Temperature, °C 300-400
Catalyst material Proprietary
removal efficiency, %a 90.0
a. [(Inlet-outlet) /inlet] 100.
Kurabo Knorca Process —
The Kurabo Knorca process is also a dry, NOx-only, SCR process.
However, in contrast to the other N0x-only FGT processes, the Knorca process
has four reactor trains in a 500-MW coal-fired power plant. Also, a moving-
bed reactor is used. The NH3 injected Into the flue gas is not diluted, but
rather, static gas mixers are included in the flue gas ducts to ensure
adequate mixing of the gases. The design conditions and removal efficiency
in the reactor are shown in Table 14. The overall design system includes a
limestone FGD system and a hot ESP.
UOP SFGT-N Process—
The UOP SFGT-N dry, NOx-only, SCR process consists of two processing
trains in a 500-MW coal-fired application with a parallel passage, fixed-bed
reactor. The design conditions and removal efficiencies in the catalytic
reactors are shown in Table 15. In order to ensure good mixing and flow
15
-------
TABLE 14. DESIGN CONDITIONS AND REMOVAL EFFICIENCIES
FOR THE CATALYTIC REACTOR IN THE KNORCA NOY-ONLY PROCESS
A
_ Item _ Reactor
Type Moving -bed
Operating conditions
Pressure drop, in. of H20 4
Space velocity, ft3/hr/ft2 7,500
NH3:NOX mol ratio 0.9:1.0
Temperature, °C 400
Catalyst material Fe-oxide on Ti02
NOX removal efficiency, %a 90.0
a. [(Inlet-outlet) /inlet] 100.
TABLE 15. DESIGN CONDITIONS AND REMOVAL EFFICIENCIES F, OR THE
CATALYTIC REACTOR IN THE UOP SFGT-N, NOX-ONLY PROCESS
__ _ Item _ Reactor _
Type Parallel passage
Operating conditions
Pressure drop, in. of H20 5.0-6.0
Space velocity, ft3/hr/ft2 7,500
NH3:NOx mol ratio 1.1:1.0
Temperature, °C 400
Catalyst material CuO and CuS04 on alumina
removal efficiency, %a 90.0
a. [(Inlet-outlet) /inlet] 100.
control and because of inflammability limits, the NH^ injected is first
diluted with air to 1.0 weight percent (5% by volume) NH3. The overall
process design for the UOP SFGT-N system includes a limestone FGD system
and a high-efficiency cold ESP.
UOP SFGT-SN Process—
The UOP SFGT-SN process is the only dry, simultaneous S0x-N0x, SCR
process included in this study and uses a parallel passage, fixed-bed reactor.
With the exception of the eight reactors, one train of processing equipment
is used in a 500-MW coal-fired application. The design conditions and removal
efficiencies in the reactors are shown in Table 16.
16
-------
TABLE 16. DESIGN CONDITIONS AND REMOVAL EFFICIENCIES FOR
THE REACTOR IN THE UOP SFGT-SN, SOx-NOx PROCESS
_ Item _ Reactor
Type Parallel passage
Operating conditions
Pressure drop, in. of H20 15.0-16.0
Space velocity, ft^/hr/ft2 7,500
NH3:NOX mol ratio 1.1:1.0
Temperature, °C 400
S02 sorbent material CuO on alumina
NOX catalyst material CuS04 on alumina
Removal efficiency, %a
SOo 90
70
T. [(Inlet-outlet) /inletj 100.
The process design includes the equipment required to convert the
absorbed S02 to 98% HaSOA. Normally the ESP would be considered
as part of the boiler cost; however, the cost of the ESP is included in
this study in order to compare the process with wet FGT systems, which
have a prescrubber for particulate removal.
Limestone FGD —
The limestone FGD system with four trains of absorbers is designed for
90% removal of S02 only; it includes neither an ESP nor a prescrubber. Each
of these scrubbers is a Turbulent Contact Absorber (TCA) type with a
presaturator for cooling and humidifying the flue gas. The design removal
efficiencies for both the presaturator and the absorber are given in Table 17.
Each scrubber is equipped with a chevron-type entrainment separator at the
scrubber outlet. Stack gas reheat to 175°F for plume buoyancy is provided
by indirect steam reheat of the cleaned gas.
Waste sludge is disposed of by ponding as a 15 weight percent solids
slurry. The settled sludge contains 60% free water; the excess water is
returned to the scrubber system.
ESP--
Because of the various ESP locations and PM removal efficiencies
required with the different NOX FGT processes, four different ESP cases
are included in this study. With three of the dry NOX FGT processes, i.e.,
Hitachi Zosen, UOP SFGT-N, and UOP SFGT-SN, a cold ESP is used which is
located just downstream of the air heaters. Four trains of ESP are used.
With the UOP SFGT-N and Hitachi Zosen processes, the ESP units are sized
and coats assigned on the basis of a 99.5% PM removal efficiency and a
specific collecting area of 300 ft2 of collecting electrode per kft3/min of
flue gas. With the UOP SFGT-SN process, since some 803 is also removed,
17
-------
removing fly ash is more difficult. For 99.5% PM removal efficiency, the
units are sized and costs assigned on the basis of a specific collecting
area of 460 ft^ of collecting electrode per kft-Vmin of flue gas. With the
Kurabo Knorca process, the remaining dry NOX FGT process, a hot ESP upstream
of the NOX reduction reactor is used. Four ESP units are used. These units
are sized and costs assigned on the basis of a 98% PM removal efficiency and
specific collecting area of 240 ft^ of collecting electrode per kft^/min of
flue gas. The remaining PM removal is achieved by the NOX reduction reactor
for a total of 99.5% removal efficiency. The only wet NOX FGT process using
an ESP is the Moretana Calcium process. A PM removal efficiency of 99.5% is
achieved by using a 94% efficient cold ESP followed by a prescrubber. A
single ESP is included. The ESP is sized and costs are assigned on the basis
of a specific collecting area of 150 ft2 of collecting electrode per kft^/min
of flue gas.
TABLE 17. DESIGN OPERATING CONDITIONS AND REMOVAL EFFICIENCIES
FOR THE PRESATURATOR AND THE ABSORBER IN THE LIMESTONE FGD SYSTEM
Item Presaturator Absorber
Operating conditions
Pressure drop, in. of H20 -a
L/G,
gal/kaft3
4.0
8.0
50.0
Removal efficiency, %b
so2
S03
HC1
NOX
5
50
100
0
90
50
-
0
a. Included in the pressure drop for the absorber.
b. Defined as [(inlet-outlet)/inlet]100.
Raw Materials—
All raw materials used in the eight NOX FGT systems are assumed to be
received by either rail or truck. Thirty-day storage facilities are provided
for them. Feed bins and intermediate process tanks are designed for an 8-
hour storage capacity.
Listed below are the primary raw materials and their typical properties.
1. Limestone (Asahi, IHI, Moretana Calcium, and limestone FGD)
Analysis: 90% calcium carbonate
2. Na4-EDTA (Asahi)
Analysis: 39% aqueous solution of Na^-EDTA
3. Soda ash (Asahi)
Analysis: 98% Na2C03
4. Sulfuric acid (Asahi, IHI) ,
Analysis: 98% H2S04
18
-------
5. Caustic soda (IHI)
Analysis: 50% NaOH
6. Sodium chloride (IHI)
Analysis: 97% NaCl
7. Slaked lime (Moretana Calcium)
Analysis: 93% Ca(OH)2
8. Ammonium sulfate (Moretana Calcium)
Analysis: 99%
9. Hydrogen chloride (Moretana Calcium)
Analysis: 35% HC1
10. Ammonia (Hitachi Zosen, Kurabo, UOP)
Analysis: Guaranteed 99.5% NH3, 82.2% N
11. Naphtha (UOP SFGT-SN)
Analysis: 133,000 Btu/gal
ECONOMIC PREMISES
The following criteria provide a common basis for the economic evalua-
tion of each of the eight alternative NOx FGT processes and make the results
directly comparable.
.Capital Investment
•Capital investment estimates are based on a midwestern plant location
and represent projects beginning in mid-1977 and ending in mid-1980, with an
average cost basis for scaling of mid-1979. Direct investments are calcu-
lated using the average annual Chemical Engineering cost indices for the
period up to 1975 and the recent TVA projections for the period 1976-1981
shown in Table 18. Actual equipment cost estimates are based on current
(1978) cost information .obtained from engineering-contracting, processing,
and equipment companies.
For the dry FGT systems, only the incremental investment and operating
cost for the larger ID fans required is charged to the FGT system. This
incremental cost is estimated as the difference between the larger ID fan
required for the combined boiler-FGT system pressure drop and the ID fan
which would be needed if no NC^ removal facilities were installed. This
option is usually less expensive than the addition of a separate fan to
compensate for the pressure drop through the NOx reactor.
Costs related to equipment, material, and construction-labor shortages
with accompanying overtime pay incentive and costs for the generation
facilities for electricity used by the NOx removal processes and FGD unit
are not included.
19
-------
TABLE 18. COST INDICES AND PROJECTIONS
Year
Plant
Material5
Labor0
1970
125.7
123.8
137.4
1971
132.3
130.4
146.2
1972
137.2
135.4
152.2
1973
144.1
141.9
157.9
1974
165.4
171.2
163.3
1975
182.4
194.7
168.6
1976a
197.9
210.3
183.8
1977a
214.7
227.1
200.3
1978a
232.9
245.3
218.3
1979a
251.5
264.9
237.9
1980a
271.6
286.1
259.3
1981a
293.3
309.0
282.6
a. Projections.
b. Same as index in Chemical Engineering for "equipment, machinery, supports.'
c. Same as index in Chemical Engineering for "construction labor."
-------
Direct Investment—
The purchase cost of the process equipment and the cost of materials
and labor for the installation of this equipment are included. Also, the
cost of piping and insulation, ductwork, excavation, site preparation,
foundation, structural, roads and railroads, electrical, instrumentation,
buildings, painting, services, and pond construction required for each unit
area are estimated. The costs for the NOX FGT processes and ESP's are
estimated based on the flow diagram, equipment lists, and TVA experience.
The calculation of the direct investment for the limestone FGD system, on
the other hand, is based on detailed engineering design.
Services, utilities, and miscellaneous costs are estimated at 6% of the
total direct investment excluding pond construction cost. This expense
covers allocated costs for the use of such power plant facilities as mainte-
nance shops, stores, communications, security, and offices. Parking lots,
walkways, landscaping, fencing, and vehicles are also included in the
service facility estimate.
Indirect Investment—
The indirect investment includes costs for engineering design and
supervision, architect and engineering contractor expenses, construction
expenses, contractor fees, and contingency. The engineering design and
supervision and contingency factors are based on proven design, not
first-of-a-kind installation.
. Engineering design and supervisloji—This Indirect investment factor
was estimated using a technique that correlates the number of major pieces
of process equipment with drafting room man-hour and engineering design
costs. Pond engineering and design costs were determined1 by a separate
procedure based on pond construction expense. In a similar manner, the
engineering design and supervision costs for landfill disposal cases are
based on major items excluding earthmoving equipment. The sum of these
costs, where applicable, appears in the indirect investment display as
engineering design and supervision for each process.
Architect and engineering (A&E) contractor expenses—This cost is
based directly upon the engineering design and supervision costs. A&E
expense is assumed to be 25% of the portion of engineering design and
supervision costs associated with major equipment. For cases involving
disposal panda, 10% of the engineering design and construction expenses
associated with pond construction is estimated as the additional A&E
expense. For the landfill systems, 25% of the total engineering design
and supervision expense is assumed to be the A&E expense.
Construction expense—Construction expense is expressed as a function
of the direct investment and was estimated by the following equation.
21
-------
Construction expense = 0.25 (x)°-83 + 0.13 (y)0-83
where x = direct investment, excluding pond or landfill
investment cost, in M$
y = direct pond cost in M$
Construction facilities (which include costs for mobile equipment, temporary
lighting, construction roads, raw water supply, safety and sanitary
facilities, and other similar expenses incurred during construction) are
considered a part of construction expenses.
Contractor fees—The relationship between contractor fees and total
direct investment used to estimate the cost of contractor fees was:
Contractor fees = 0.096 (a)0-76
where a = direct investment in M$
Contingency—The contingency is assumed to be 20% of the sum of total
direct investment, engineering design and supervision costs, A&E expenses,
construction expenses, and contractor fees.
Other Capital Charges—
The sum of the total direct and total indirect investment is total
fixed investment. Other capital charges which must be added to this total
fixed investment to obtain the total capital investment are: (1) allowance
for startup and modifications, (2) interest during construction, (3) land,
(4) working capital, and (5) royalty fees.
Allowance for^ startup and modifications—This expense is estimated to
be 10% of the total fixed investment excluding pond construction.
Interest during construction—This item was estimated to be 12% of the
total fixed investment. This percentage is calculated as the simple interest
which would be accumulated at a 10%/yr rate assuming a debt-equity ratio
of 60:40 on the incremental capital investment and a 3-year project expendi-
ture schedule as shown in Table 19.
Land—The cost of land is estimated at $3500/acre.
Working capital—Working capital consists of: (1) money invested in
raw materials, supplies, and finished products carried in stock and semi-
finished products in the process of being manufactured; (2) accounts
receivable; (3) cash retained for payment of operating expenses, such as
salaries, wages, and raw material purchases; (4) accounts payable; and
(5) taxes payable (14). For these premises, working capital is defined as
the equivalent cost of 3 weeks of raw material costs, 7 weeks of direct
costs, and 7 weeks of overhead costs.
Royalty fees—The royalty fees were added based on information supplied
by the vendor. If there were no data, a $0.3M fee was assumed. For both UOP
processes, the royalty fee was prorated for inclusion in the direct investment.
22
-------
TABLE 19. PROJECT EXPENDITURE SCHEDULE
Year
Total
Traction of total expenditure
as borrowed funds 0.15 0.30 0.15 0.60
Simple interest at 10%/yr as
% of total expenditure
Year 1 debt 1.5 1.5 1.5 4.5
Year 2 debt - 3.0 3.0 6.0
Year 3 debt _z _i 1*5 1.5
Accumulated interest as % of
total expenditure 1.5 4.5 6.0 12.0
Revenue Requirements
Annual revenue requirement calculations are based on 7000 hours of
operation per year using mid-1980 operating costs and average capital
charges. Process operation schedules are assumed to be the same as the
power plant operating profiles and remaining life assumptions given in the
power plant design premises.
Direct Costs—
Direct operating costs include raw material, labor, utility, and
maintenance costs. The 1980 projected unit costs for the raw materials
used-in the various NOx FGT processes are given in Table 20 as delivered
costs to a midwestern location. Because of the proprietary nature of the
catalyst used in some of these FGT procesi.es, the costs for these catalyst
are listed separately and are given only as annual quantities in Table 20.
The catalyst cost is based on vendor quotes and a guaranteed life of 1 year.
The projected labor costs assumed for 1980 are shown in Table 21.
Projected 1980 utility costs are shown in Table 22. Unit costs for
electricity and steam generated by the power plant are based on actual
production cost including labor, fuel, depreciation, taxes, and rate base
return on investment. The electricity rates are based on purchase from an
independent source with full capital recovery provided and are adjusted for
the quantity used. Process water rates vary with the quantity of water
required. If a heat credit is applied to a process, $2/MBtu is used as a
basis for the credit.
Maintenance costs are estimated as a percentage of the direct invest-
ment, excluding pond construction. The maintenance factors vary according.
to operating characteristics of the process and are estimated based on
either actual operating experience or maintenance needs in similar process
areas. The estimated maintenance factors for each process are shown in
23
-------
TABLE 20. PROJECTED 1980 COSTS FOR RAW MATERIALS
Raw material
Limestone
Na2C03
Na4 • EDTA
H2S04
K2S04
Copperas
Slaked lime
CuCl2
HC1
(NH4)2S04
Powdered limestone
NaCl
NaOH
NH3
Naphtha
Cost, $/ton
7.00
103.00
743.00
54.00
144.00
86.50
60.00
2,105.00
81.00
82.00
32.00
25.80
187.00
150.00
285. 60a
Catalyst Cost, $/yr
Moretana Calcium 4,116,200
Hitachi Zosen 6,880,000
Kurabo 4,316,000
UOP SFGT-SN 6,648,000
UOP SFGT-N 2,449,000
a. Based on $0.50/gal with 42 gal/bbl and 13.6
bbl/ton.
TABLE 21. PROJECTED 1980 UNIT COST FOR LABOR
Type .Unit cost, $/man-hr
Plant operation and supervision 12.50
Analyses 17.00
Landfill equipment and truck
operators 17.00
Table 23. Pond maintenance is calculated separately as 3% of the pond
construction cost. The maintenance charge for the ozone system in the
IHI process was assumed to be 4% of ozone plant investment. Maintenance
for the landfill section was also assumed to be 4% of the total investment
for the landfill.
24
-------
TABLE 22. PROJECTED 1980 UNIT COST FOR UTILITIES
Utility Cost, $
Steam 2.00/MBtu
Electricity 0.029/kWh
Process water (UOP) 0.06/kgal
Process water 0.12/kgal
Naphtha 0.50/gal
TABLE 23. ESTIMATED ANNUAL MAINTENANCE FACTOR
Percentage of
Process direct investment
Asahi 7
Hitachi Zosen 4
IHI ?a
Kurabo 4^
Moretana Calcium 7
UOP SFGT-SN 4
UOP SFGT-N 4
Limestone FGD 8b
Maintenance of ozone section of IHI is 4%
of investment for ozone manufacturing
plant.
Excluding pond construction. Pond mainte-
nance is estimated as 3% of pond construc-
tion cost and maintenance for landfill is
assumed to be 4% of investment.
Other direct costs such as land preparation cost (i.e., the annual
expense for clearing, excavating, draining, and reclaiming the disposal
site) are based on the annual quantity of land used, and are assumed to
be $1700/acre. The operating cost of trucks and landfill equipment (fuel
and maintenance) for a 1-mile travel distance is estimated at $0.06/ton
of wet sludge. A fuel and maintenance cost of $C.16/ton of wet sludge is
assumed for earthmoving equipment used in the landfill operation (2).
Indirect Costs—
Indirect costs Include capital charges, projected to 1980, and overheads.
Following power industry practice, regulated company economics are used for
25
-------
calculating the capital charges. A breakdown of the capital charges is
given in Table 24. The depreciation rate is straight line based on the life
of the power plant.
TABLE 24. ANNUAL CAPITAL CHARGES FOR POWER INDUSTRY FINANCING
Percentage of
total depreciable
capital investment
Years remaining life 30
Depreciation-straight line (based on
years remaining life of power unit) 3.3
Interim replacements (equipment having
less than 30-yr life) 0.7
Insurance and property taxes 2.0
Total rate applied to original 6.0
investment
Percentage of
unrecovered
capital investment5
Cost of capital (capital structure
assumed to be 60% debt and 40% equity)
Bonds at 10% interest
Equity*5 at 14% return to stockholder
Income taxes (Federal and State)c
Total rate applied to depreciation 17.2^
base
a. Original investment yet to be recovered or "written off."
b. Contains retained earnings and dividends.
c. Since income taxes are approximately 50% of gross return,
the amount of taxes is the same as the return on equity.
d. Applied on an average basis, the total annual percentage
of original fixed investment for new (30 yr) plants would
be 6.0% + 1/2 (17.2%) = 14.6%.
In estimating the regulated capital charges associated with FGT, the
conventional method of considering the overall life of the power plant is
used. FPC recognizes the conclusion of the National Power Survey that a
30-year service life is reasonable for steam-electric plants. Because some
life spans are less than 30 years, however, FPC has designated interim
replacements as an allowance factor to be used in estimating annual revenue
26
-------
requirements. Use of this allowance following FPC-recoiranended practice
provides for financing the cost of replacing such short-lived units. An
average allowance of about 0.35% of the total investment is normally provided.
However, to provide for the unknown life span of S0x-N0x control facilities,
a larger allowance factor of 0.70% is used for new units. An insurance
allowance of 0.5% is also included in the capital charges based on FPC
practice. Property taxes are estimated at 1.5% of the total depreciable
capital investment.
Debt-equity ratio is another component of capital charges for which
variations of ratios may be expected. FPC data indicate that the long-term
debt for privately owned electric utilities varied only slightly from 51.5
to 54.8% of total capitalization during the period 1965-1973. However,
recent economic trends have changed the incremental debt-equity ratio
because utilities are more dependent on bonds and bank loans for project
funding. For this study, the capital structure is assumed to be 60% debt
and 40% equity. The interest rate for bonds is assumed to be 10% and the
return to stockholders on equity 14%. Costs of capital and income tax
charges are applied to the uncovered portion of capital investment. Income
taxes and return on equity are each about 50% of gross return. Since return
on equity is 5.6% of total capital investment, income taxes are also 5.6% of
total capital investment (see Table 24). Since most, regulatory commissions
base the annual permissible return on investment on the remaining deprecia-
tion base (that portion of the original investment yet to be recovered or
"written off"), a portion of the annual capital charge included in the
lifetime operating costs declines uniformly over the life of the power plant.
Plant, administrative, and marketing overheads vary from company to
company. Based on the various methods used in industry and illustrated in
a variety of cost estimating sources, the following method of estimating
overheads is used. Plant overhead is estimated as 50% of the subtotal
conversion costs less utilities, which include the projected costs for
labor, maintenance, and analyses. Administrative overhead is estimated as
10% of operating labor and supervision. Marketing the byproduct is also a
consideration. Marketing is estimated as 10% of sales revenue.
Spent Catalyst Disposal-r
For the dry FGT processes, the cost of disposing of the spent catalyst
is added to the annual revenue requirement. For the Kurabo Knorca, UOP SFGT-
N, and UOP SFGT-SN processes, disposal cost for the catalyst carriers (Ti02,
alumina, alumina, respectively) is used as the basis for the yearly disposal
charge. For the Hitachi Zosen process, the catalyst base support has scrap
metal value. Thus, this scrap value yields a credit toward the annual revenue
requirement.
Byproduct Sales—
In the evaluation of annual and lifetime economics, credit from sale of
byproducts is deducted from the yearly operating cost to give the net effect
of the FGT process on the cost of power. Since byproduct credits are
provided only for sulfuric acid in this study, the UOP SFGT-SN simultaneous
S0x-N0x control process is the only system given a byproduct credit. A 1980
projected sales price for the byproduct sulfuric acid from the UOP process
was estimated at $30/ton.
27
-------
SYSTEMS ESTIMATED
Process descriptions, flowsheets, material balances, and major equip-
ment lists and descriptions were prepared for each of the eight FGT processes
included in this evaluation. Each process was divided into major operational
areas to facilitate investment and revenue requirement comparisons for
similar processing steps. Using the material balances, each piece of equip-
ment was described, sized, costed, and listed by processing subsection, with
total equipment costs in 1979 dollars for each item. The additional items
of cost in each area which were not shown in the equipment lists, (i.e.,
installation labor and material costs for electrical, piping, ductwork,
foundations, structural, instrumentation, insulation, and site preparation)
were incorporated in the area investment estimates shown in the summary
capital investment estimates in the following section.
CLASSIFICATION OF NOX REMOVAL PROCESSES
The NOx FGT processes can be divided into two types, dry or wet, based
on whether or not the NOx is absorbed into an aqueous solution. The primary
methods of dry NOx FGT which have been studied are: (1) catalytic decom-
position, (2) SCR, (3) nonselective catalytic reduction, (4) selective
noncatalytic reduction, (5) adsorption, and (6) electron beam radiation (6).
About half of all the NOx FGT processes currently undergoing development is
the SCR type, which is the most widely tested and technically advanced type
of FGT process. SCR is the only type of dry NOX FGT process included in
this study.
All current SCR-type processes use NH3 to selectively reduce NOX to
molecular nitrogen and water in the presence of a catalyst. The reactions
of NH3 with NOx are usually expressed by the process developers as one of
the following:
4NH3(g) + 6NO(g) * 5N2(g) + 6H20(g) (1)
8NH3(g) + 6N02(g) - 7N2(g) + 12H20(g) (2)
4NH3(g) + 4NO(g) + 02(g) * 4N2(g) + 6H20(g) (3)
4NH3(g) + 2N02(g) + 02(g) + »2(g) + 6H20(g) (4)
Reactions (3) and (4) may more accurately represent the reduction reactions
since several studies have demonstrated that the presence of at least some
28
-------
oxygen improves the NOX- reduction efficiency (3, 4, 9, 10). As 90-95% of the
total NO formed during the combustion of coal is present in the flue gas in
the form of NO and N02 is only 5-10%, reaction (3) was used as the basis for
material balance calculations in this study.
Process developers reported that most of the excess NH3 reacts with
oxygen as follows:
4NH3(6) + 3°2(g) * 2"2(8) + 6H2°(g)
The NH3 level in the treated flue gas leaving the FGT systems is
low, reportedly less than 20 pptn.
There are some important factors affecting performance of these SCR
processes which should be mentioned. The operating temperature range for
most of these processes is about 300-400°C, though the more efficient NOX
removal usually occurs in the higher portion of this range. To maintain the
reactor temperature at desirable operating levels during periods of reduced
boiler load, most process vendors recommend bypassing a part of the flue gas
around the economizer. In some pilot-plant and larger operations, auxiliary
heaters have been used to maintain reactor temperatures during turndown.
However, if this were necessary for large utility boilers, the economics of
the SCR processes would be significantly increased.
The formation of ammonium bisulfate [(NH4)HS04] and ammonium sulfate
[(NH£)2S04] is another problem to consider. (NH4)HS04 condensation usually
occurs downstream of the reactor as the flue gases are cooled in the air
heater. This presents potential corrosion problems and also reduced heat
transfer problems. Also, some vendors have indicated problems with deposits
of (NH4)2S04 in the reactor. Soot blowers for the reactor and air heater
are the most commonly reported treatment of this situation. However, the
information about this problem is still limited and the reported frequency
of use of the soot blowers or other means varies greatly among vendors.
Thus, an adequate evaluation of this problem is not possible here.
Another problem which may need to be addressed in future development
programs is the monitoring of low concentrations of NH3 and NOX in the
reactor outlet. Most vendors state that the NH3 feed rate is partly
controlled based on the NOX and NH3 concentrations in the flue gas leaving
the reactor. There are fears by some, however, that monitoring of these
low concentrations of NOX and NH3 may be difficult and that (NH4)HSC>4
or (NH4)2S04 deposits in the sampling lines and instrument probes may
compound the problem.
The dry SCR processes included in this report are offered by Hitachi
Zosen; Kurabo Industries, Ltd.; and UOP, Inc. (UOP offers both the simul-
taneous SOX-NOX process and the N0x-only process) .
29
-------
The wet FGT processes were, In most cases, designed to take advantage
of technology already available from the previously developed FGD systems.
Most of the wet processes were originally designed as simultaneous SOx-NOx
systems. Unfortunately, the solubility of NO, composing 90-95% of the total
NOx present in coal-fired boiler flue gas, is low in aqueous solution and
thus imposes a formidable obstacle to wet NOX FGT processes. N02» which is
the lesser component of the NOX, is much more soluble than NO and does not
impose as difficult a problem, although the solubility of N02 is also poor
relative to SOX. Therefore, the major task associated with any wet NOX
removal process is the absorption of the NOX by the scrubbing solution where
it can be concentrated and more economically converted into other forms.
There are two common methods of removing the NOx from the flue gas:
(1) direct absorption of the NOX in the absorbing solution or (2) gas-phase
oxidation to convert the relatively insoluble NO, either partially or
entirely, to N02« Thus, the wet NOX removal processes can be classified
as absorption or oxidation processes, depending on whether or not the flue
gas is treated with a gas-phase oxidant before absorption.
Additionally, each of these classifications can be divided based on the
method by which the NOX is converted after it has been absorbed by the
scrubbing solution. Processes which reduce the absorbed NOx either partially
or completely to molecular nitrogen or complex nitrogen-sulfur compounds
can be classified as reduction processes. Processes which do not reduce the
absorbed NOx are absorption processes. Thus, the wet NOX processes can be
categorized into one of the following groups: oxidation-absorption-reduction,
oxidation-absorption, absorption-reduction, or absorption-oxidation. Figure 1
shows this classification system.
Because of the limited development status, technical feasibility, and
also the doubtful economics of some of the process classifications, only two
classifications of these wet NOx FGT processes are evaluated in this report:
oxidation-absorption-reduction and absorption-reduction processes. Included
in the oxidation-absorpticn-reduction category is the Moretana Calcium process
offered by Sumitomo Metal Industries and IHI process. The Asahi Chemical
process is the example of absorption-reduction technology included in this
study.
The oxidation-absorption-reduction processes are based on the use of a
gas-phase oxidant to convert essentially all of the NO to N02« Before
entering the absorber the flue gas is injected with an oxidant agent, either
ozone (used by IHI) or C102 (used in Moretana Calcium process).
This gas-phase oxidant rapidly and selectively oxidizes the NO to N02
which is then absorbed into an aqueous sulfite solution. The SOX in the
flue gas is simultaneously absorbed and forms either the sulfite or bisul-
fite ion In the scrubbing solution. The sulfite ions are partially oxidized
to sulfate during reduction of the absorbed N02 and, depending on the process,
the remaining sulfite may be oxidized by air to sulfate in the regeneration
section. The sulfate is removed as gypsum. The absorbed N02 is at least
partially reduced and removed as a combination of nitrate salts, molecular
nitrogen, and complex nitrogen-sulfur compounds.
30
-------
INLET
I FLUE GAS
CONTAI MING
1 NO, j
YES
ARE
NITROGEN
COMPOUNDS
REDUCED
GAS PHAS
OXIDATION
ARE
NITROGEN
COMPOUNDS
». OXIDATION-ABSORPTION-REDUCTION
OXIDATION- ABSORPTION
ABSORPTION - REDUCTION
ABSORPTION-OXIDATION
Figure 1. Classification system for wet NOx removal processes,
-------
Because the Moretana Calcium process uses CIC^ to oxidize NO to NC>2,
approximately half of the absorbed N0£ is reduced to molecular nitrogen and
the other half is converted to nitrate salts. The IHI process, which uses
ozone as the gas-phase oxidant, is based on limestone slurry scrubbing with
a proprietary catalyst added to the scrubbing solution to enhance NC^
absorption. This use of ozone and the addition of the catalyst apparently
results in most ~>f the N02 being reduced to complex nitrogen-sulfur compounds
such as sulfamine salts. Only relatively minor amounts of molecular nitrogen
and nitrate salts are formed in the process. These complex nitrogen-sulfur
compounds are then thermally decomposed into molecular nitrogen and sulfate
salts.
The absorption-reduction processes were specifically developed for the
simultaneous removal of SOX and NOX from power plant flue gas without using
an expensive gas-phase oxidant. These processes are based on certain ferrous-
chelating compounds which improve the absorption of the relatively insoluble
NO. The NO is absorbed by the scrubbing solution and forms a complex with
the chelating compound. The SOX is simultaneously absorbed by the scrubbing
solution as the sulfite ion and reacts with the NO complex. The NO is
reduced to molecular nitrogen and the ferrous-chelating compound is regenerated
while the sulfite ion is being oxidized to sulfate. This sulfate is generally
removed as gypsum.
ASAHI CHEMICAL SOX-NOX PROCESS - WET, ABSORPTION-REDUCTION
Process Description
This process developed by the Asahi Chemical Industry Company uses sodium
sulfite (Na2S03) solution to absorb both SOX. and NOX (16). A ferrous-chelating
compound is used to aid in the absorption of NOX. The five main sections of
this system are: (1) prescrubbing, (2) absorption of both SOX and NOX,
(3) reduction of NOX, (4) decomposition of dithionate, and (5) gypsum produc-
tion. The Asahi process, as a whole, is an intricate system made up of many
simple processing units.
The flue gas from the air heater is passed through a prescrubber which
removes 99% of both particulates and chlorides. In addition, some of the
scrubbing solution evaporates to humidify and to adiabatically cool the flue
gas from 300°F to 127°F. The liquid prescrubber effluent is recirculated,
makeup water is added, and small purge stream is removed and pumped to a fly
ash thickener. The thickener overflow stream is recycled to the prescrubber
and the bottoms containing the fly ash and the chlorides are removed, neutral-
ized by limestone, and pumped to a waste disposal pond. Thus, the fly ash
and chlorides in the flue gas are removed and treated before the flue gas
reaches the absorber.
After leaving the prescrubber the flue gas enters a packed-bed absorber
where it flows countercurrent to a 6.3 pH sodium-salt scrubbing solution
containing ferrous ethylenediaminetetraacetic acid (Fe+2-EDTA). As the flue
gas passes through the lower portion of the absorber, the S02 is rapidly
absorbed into the solution and undergoes the following reactions:
32
-------
S02(g) * S°2(aq)
Na2S°3(aq) + S°2(aq) + H2° " 2NaHS03(aq)
2NaHS03(aq) + l/202(aq) + Na2S206 + H20 (3)
The NOX, on the other hand, is gradually absorbed over the entire length
of the absorber and undergoes the following reactions to form f errous-chelate
complex:
N°(g) * N°(aq)
N°(aq) + Fe'EDTA " FeEDTA'NO (5)
In addition to these primary reactions, other side reactions occur
simultaneously. Both the Fe+2-EDTA and the Na2503 are readily oxidized
In the scrubbing solution by oxygen absorbed from the flue gas to form
ferric EDTA (Fe+3*EDTA) and Na2SC>4:
°2(g) * °2(aq)
Na2S°3(aq) + 1/2°2(aq)
Fe+2-EDTA(aq) °? Fe+
The flue gas leaving the absorber is passed through a mist eliminator,
reheated by mixing with the flue gas from the thermal decomposers, and sent
to the stack.
The liquid effluent from the absorber is pumped to a reducing tank
where makeup soda ash (Na2C03) is added to replace sodium lost from the system
and to convert sodium bisulfite (NaHSOs) to Na2S03, The reactions occurring
in the reducing tank include;
Fe+2.EDTA.NO(aq) + Na2S03(aq) - Fe+2.EDTA(aq) + Na2S04(aq) + l/2N2(g)
Na2C03(aq) + 2NaHS03(aq) + 2Na2S03(aq) + H20 + C02(g) (10)
Most of the NO absorbed is thus converted to molecular nitrogen and
vented to the atmosphere. In addition, the Fe+3. EDTA which was produced in
the absorption column (see equation 8) is reduced by the sulfite ion to
Fe+2.EDTA.
Most of the effluent from the reducing tank is recycled to the absorber.
The remainder (10-20% of the circulating solution) is pumped to a multiple-
33
-------
effect evaporator system in the regeneration section. The concentrated
solution from the evaporators is then pumped to a cooling crystallizer where
hydrated sodium dithionate (Na2S20g) and Na2S04 crystals are produced under
vacuum at 50°F. These crystals are separated from the mother liquor in a
screw decanter and sent to a dryer operating at 250-300°F in which the
hydrated crystals are converted to anhydrous sodium salts. For example, the
reaction for the Na2S206 crystals is:
Na2S206-2H20(s) £ Na2S206(s) + 2H20(g) (11)
Most of the mother liquor from the decanter is recycled to the reducing
tank and a smaller stream is passed through a sodium imidodisulfonate
[NH(S03Na)2] treatment section. This section is necessary since soluble
NH(S03Na)2 is formed in small quantities in the absorber and must be removed
to prevent its buildup in the scrubbing solution. The NH(S03Na)2 is con-
verted to relatively insoluble potassium imidodisulfonate [NH(S03K)2] by
reaction with potassium sulf ate (K2S04) :
NH(S03Na)2(aq) + K2S04(aq) + NH(S03K)2(s) + Na2S04(aq) (12)
The crystals of NH(S03K)2 are separated in a screw decanter and sent
to a thermal cracker. The mother liquor, mostly dissolved Na2S04, is
recycled to the reducing tank. The thermal cracker, which is heated
indirectly by flue gas from an oil-fired furnace, operates at about 930°F
and decomposes the NH(S03K)2 into S02, molecular nitrogen, potassium sulfite
(K2S03) , and K2S04 by the following reaction:
2NH(S03K)2(s) I 2S02(g) + K2S04(s) + K2S03(||) + H20(g) + N2(g) (13)
The concentrated S02 gas stream is sent to the Na2SC»4 converter
while the solid K2S03 and K2SO, are recycled to the imidodisulfonate
treatment section.
The anhydrous Na2S2Og and Na2SO^ mixture from the dryer is sent to a
thermal cracker where it is heated indirectly to about 572°F to decompose
the Na2S2Og by the following reaction:
Na2S206(s) I Na2S04(g) + S02(g) (14)
A portion of this Na2S04 is removed as a byproduct, but most is sent
to a premixing tank where it is dissolved in makeup water and reacted
with calcium sulfite (CaS03) to produce a slurry. This slurry is pumped to
a Na2S03 converter where it is treated with the S02 from both the dithionate
and the imidodisulfonate thermal crackers to form gypsum and NaHSO., by the
following reactions:
-------
CaS03(s) + S02(g) + H20 -> Ca(HS03)(aq) (15)
Ca(HS03)2(a x + Na2S04(aq) + 2H20 -> 2NaHS03(aq) + CaS04.2H20(g) (16)
The converter product stream is pumped to a screw decanter where gyspura
is removed as a solid byproduct. The liquor is sent to an S02 stripper,
operating at 200°F, pH 4, and a pressure of 100 mmHg, where the bisulfite
salts are partially converted to sulfites and S02:
Ca(HS03)2(aq) + CaS03-l/2H20(s)* + S02(g)t + 1/2H20 (17)
2NaHS03(aq) + Na2S03(aq) + S02(g)+ + H20 (18)
The S02 produced in the stripper is recycled to the sodium sulfate
converter and the effluent containing the sulfite is sent to the double
decomposition reactor where the Na2S03 is converted to CaS03 by the
following reaction with makeup limestone slurry:
2NaHS°3(aq) + CaC°3(s) *
Na2S03(a s + CaS03-l/2H20(g)4- + 1/2H20 + C02(g) (19)
The effluent from the double decomposition reactor is pumped to a thick-
ener where the bottoms, a CaSO^ slurry, are removed and pumped to a screw
decanter. The solid from the decanter is then reslurried with makeup
water and recycled to the sodium sulfate converter. The liquor from the
decanter is recycled to the gypsum thickener. The thickener overflow, which
is essentially a Na2S03 solution, is sent to a regenerated solution tank,
where makeup Fe+2*EDTA is added, and is then recycled to the reducing tank.
Analysis of Processing Subsections
The flow diagram and the material balance for the Asahi process are
shown in Figure 2 and Table 25 respectively. To facilitate cost determina-
tions and comparisons, the Asahi process was divided into 11 processing
sections and the processing equipment was assigned to the appropriate section.
The equipment list, giving the description and installed cost by section, is
shown in Table 26.
Each of these processing sections is described in more detail below.
Material Handling —
This and the following area compose the limestone preparation section.
The material-handling section includes all equipment needed to receive the
limestone by rail or truck and to maintain a supply of limestone to the
weigh feeders. One train of processing equipment is used.
35
-------
SO, COOLER
Figure 2. Asahi SOX-NOX process. Flowsheet.
-------
TABLE 25. ASA11T. SOX-IIOX PROCESS
MATERIAL BALANCE
Description
1
2
\
1,
•>
h
H
t>
12.
Total stream, Ib/hr
Sft3/nin (60°F)
Temperature, °F
Pressure, pslg
^pacific gravity
_pH
I'nd Is solved solids, $
1
Coal to boiler
428.600
Combustion air
to air heater
4,546,200
1,005,000
80
Combustion air
to boiler
4,101,800
906,700
535
4
Gas to
economizer
4.516,100
958,000
890
Gas to air
heater
4.516.100
958.000
705
.
Description
1
—
•(
1 4"
S
f>
JL
IL
Total stream. Ib/hr
SftJ/min (60°F)
Specific gravity
pH
TJndlasolved solids, 2
Gas to
4.960.400
1.056.000
300
Effluent from
prescrubber
3.743.200
1.06
3.5
Makeup H20 to
prescrubber
T7d
Effluent from
prescrubber
holding tank
27
12,546
1.05
Recycle to
top of
prescrubber
3,895,400
27
7.415
1.05
_j_
2
^
_ft
H
III
Pressure, paiz
Com
Specific gravity
pH __..
Undisaolved solids. »
2-701.000
. 127
S.131
1.05
Thickener
overflow to
prescrubber
holding tank
2.52n.60Q —
Thickener
bottoms to
neutralization
reactor
180.400
Ash slurry
purge to
disposal pond
206.500
OT~
Flue gas from
prescrubber to
absorber
5.113.000 1
l.i.27, 9,7.5
... , 127 1
_j
2
_)
5
h
•}
H
-------
TABLE 25. (continued)
Stream No.
Description
J
2
i
4
rj
ft
7
».
9
1°
Total stream, Ib/hr
Sft^/mln (60°F)
Temperature. °F
Pressure, pals
Gpro
Specific gravity
pH
Undissolved solids, Z
21
Effluent from
reducing tank
48 84n nno
•>.2
Absorber
recycle
47,630,000
23
Triple effect
evaporator feed
1,210,000
24
Condenser feed
75 nno
25
Crystallizer
feed
1.073,900
^"
t
\
Stream No.
Description
1
2
i
i,
ri
6
7
»
-------
TABLE 25. (continued)
Stream No.
Description
I
>
•\
4
ri
h
7
H
q
in
Total stream, Ib/hr
SftJ/rain <60°y)
Temperature, °F
Pressure, psig
Gpm
Specific gravity
PH
Undissolved solids, %
41
Fuel oil to
furnace
11,450
25
42
Combustion air
to furnace
(153; excess 02)
181,000
80
43
Flue gas to
KH(S03K)2
cracker
12,790
44
Flue gas to
N«2S206
cracker
179,500
45
Flue gas to
dryer
179,500
Stream No.
Description
t
2
t
4
5
h
7
8
•»
Iff
Total stream, Ib/hr
SftJ/min (60QF)
Temperature, UF
Pressure, pstg
Gpm
Specific gravity
pH
Undissolved solids, %
46
Flue gas from
dryer
971 T(\a
47
Flue gas
from dryer
and NH(S03K>2
cracker
238,500
48
Na2S20g feed to
cracker
80,600
49
N32S04 from
cracker
64 , 000
50
Na2S04 to
indoor storage
10.400
Stream No.
Deccrlption
1
2
1
4
">
A
7
H
9
10
Total stream, Ib/hr
SftJ/min (60°F)
Temperature, °F
Pressure, pslg
Gpm
Specific gravity
PH
Undissolved solids, %
51
Wa2S04 to surge
storage
53,600
52
Makeup H20 to
premlxing tank
31.60J3
53
S02-rich
off -gas from
Ha2S206 cracker
16.600
54
S02-rlch
off-gas from
NH(S03K)2
cracker
2.575
55
Combined
cracker off-gas
19.200
Stream No.
Description
1
2
1
4
5
6
7
H
9
iU
Total stream, Ib/hr
SftVmin (60°F)
Temperature, °F
Pressure, pslg
Gpm
Specific uravlty
pH
Undissolved solids, %
56
Effluent from
premising tank
278.100
16.5
57
Converter
off-gas recycle
to premlxing
tank
2.200
58
NazSOi,
converter
effluent to
CaS04 thickener
299.800
59
Thickener
overflow to
No.l NaK303 tanV
253.500
60
Thickener
underflow to
gypsum
centrifuge
159.700
27.5
(continued)
39
-------
TABLE 25. (continued)
1
:>
i
ri
h
7
H
i
4
>
h
7
H
i)
ID
Total stream, Ib/hr
Sft3/min (600F)
Temperature, op
Pressure, pslg
Gpm
Specific gravity
PH
Undissolved solids, %
66
S02 stripper
off-gas
50 700
67
H20 condensed
from S02
stripper
off-gas
42,200
68
S02 off-gas from
condenser
8,460
69
Combined S02
off-gas streams
to Ma 2804
converter
27,610
70
Effluent from
S02 stripper
202.800
Stream No.
Description
1
>
1
'-i
r>
h
7
H
9
10
Total stream, Ib/hr
Sft3/min (60UF)
Temperature, °F
Pressure, psig
Cpm
Specific gravity
PH
Undissolved solids, %
71
Effluent from
double
decomposition
reactor
280,700
72
Thickener
overflow to
regenerated
solution tank
227,300
73
Thickener
underflow to
CaS03
centrifuge
64,800
70
74
"entrate to No. 2
CaS03 tank
11,400
75 .
Solids from
centrifuge
to Mo. 1
CaSOs tank
53,400
85
Stream No,
Description
1
j
I
•'4
•i
i)
7
H
V
10
Total stream, Ib/hr
Sft-t/min (60°F)
Temperature, °F
Pressure, psig
Gpm
Specific gravity
PH
Undissolved solids, £
76
Makeup H20
to No. 1
CaSO} tank
95.600
77
Makeup CaS03
slurry to
preraix tank
151,300
30
78
Makeup FeS04
240
79
Makeup EDTA
250
80
Regenerated
solution to
reducing tank
227.600
{continued)
-------
TABLE 25. (continued)
Stream No.
_!_
)
H
SP
i°_
Total stream, Ib/hr
SftJ/min (60°F)
Temperature, "F
Pressure, pslg
Gpm
Specific gravity
pH
Undissolved solids, h
81
Makeup
limestone
32-lQfl
Makeup HzO to
ball mill
Makeup
limestone
slurry to
limestone
slurry feed
tank
1.61 j
84
Makeup H20 to
limestone
a lurry feed
tank
85
Effluent from
limestone
slurry feed
tank
15
DpscrintlDn
1
j
1
-3-
a
-------
TABLE 26. ASAHI SOX-NOX PROCESS
BASE CASE EQUIPMENT LIST
DESCRIPTION AND COST
Area 1—Material Handling
Item
No.
Description
Total
equipment
cost. 1979 $
1. Car shaker and hoist 1
2. Car puller 1
3. Hopper, unloading 1
4. Feeder, unloading
vibrating
5. Conveyor, unloading 1
belt
6. Conveyor, unloading
incline belt
7. Unloading pit dust
collector
8. Pump, unloading pit 1
sump
9. Conveyor, storage belt 1
10. Tripper, storage 1
conveyor
11. Mobile equipment
12. Hopper, reclaim
13. Feeder, reclaim
vibrating
14. Conveyor, reclaim
belt
1
1
20 hp shaker, 7.5 hp hoist
25 hp puller with 5 hp return
16 ft dia x 10 ft straight
side height, carbon steel
1 3.5 hp
20 ft horizontal, 5 hp
1 310 ft, 50 hp
Polypropylene bag type, 2,200
ft3/min, 7.5 hp
60 gpm, 70 ft head, 5 hp
100 ft, 2 hp
30 ft3/min, 1 hp
Scraper tractor
7 ft wide, 4.25 ft high x
2 ft wide bottom, carbon
steel
1 3.5 hp
1 100 ft, 2 hp
(continued)
31,000
52,000
12,800
14,400
17,700
89,200
19,200
4,200
45,200
16,400
137,800
1,500
14,400
32,200
•42
-------
TABLE 26. (continued)
15.
16.
17.
18.
19.
Area
1.
2.
3.
4.
5.
Item
Conveyor, reclaim
incline belt
Reclaim pit dust
collector
Pump, reclaim pit
sump
Elevator, reclaim
bucket
Bin, feed
Subtotal
2 — Feed Preparation
Item
Feeder, bin weigh
Gyratory crushers
Ball mill dust
collectors
Ball mill
Tank, ball mill
No. Description
1 193 ft, 4 hp
1 Polypropylene bag type
1 60 gpm, 70 ft head, 5 hp
1 90 ft high, 75 hp
2 13 ft dia x 21 ft straight
side height, covered,
carbon steel
No. Description
2 14 ft pulley centers, 2 hp
2 75 hp
2 Polypropylene bag type,
2,200 ft3/mln, 7.5 hp
2 9.3 tons/hr, 125 hp
2 10 ft dia x 10 ft height,
Total
equipment
cost, 1979 $
53,800
19,200
4,200
86,400
33,200
684,800
Total
equipment
cost, 1979 $
39,600
119,500
38,400
320,100
26,800
product
6. Agitator, mills
product tank
5,500 gal, open top,
flakeglass-lined, carbon
steel
10 hp, resin-lined,
carbon steel
(continued)
17,500
43
-------
TABLE 26. (continued)
Item
No.
De scrip t ion
Total
equipment
cost, 1979 $
7. Pump, ball mill
product tank
8. Tank, slurry surge
9. Pump, limestone
slurry feed
10. Agitator, limestone
slurry tank
Subtotal
Centrifugal belt drive, 55
gpm, 50 ft head, 5 hp resin-
lined, carbon steel
40 ft dia x 45 ft height,
42,300 gal, open top,
resin-lined, carbon steel
Centrifugal belt drive, 220
gpm, 100 ft head, 12 hp,
resin-lined, carbon steel
75 hp, resin-lined,
carbon steel
7,800
181,700
11,100
68,400
830,900
Area 3—Gas Handling
Item
No,
Description
Total
equipment
cost, 1979 $
1. Blowers, stack gas
Subtotal
388,200 ft3/min, AP, 39.4
in. H20, 1,500 hp, carbon
steel
2,269.400
2.269.400
Area 4—Particulate Control
Item
No.
Description
Total
equipment
cost. 1979 $
1. Prescrubbers
2. Tanks, prescrubber
recycle
Venturi, 1.97 ft wide x 32.8
ft long x 32.8 ft high,
resin-lined, carbon steel
18 ft dia, 21.5 ft high,
40,900 gal, open top, resin-
lined, carbon steel
(continued)
44
703,200
147,200
-------
TABLE 26. (continued)
Item
No,
Description
Total
equipment
cost, 1979 $
3. Agitator, prescrubber 4
holding tank
4. Pump, prescrubber 6
recycle
5. Thickener, ash
Pump, thickener
overflow recycle
Pump, thickener
underflow
8. Reactor, ash
neutralization
9. Agitator, neutraliza-
tion reactor
10. Pump, ash to pond
Subtotal
15 hp, resin-lined, carbon 86,400
steel
Centrifugal belt drive, 3,200 109,500
gpm, 144 ft head, 200 hp,
resin-lined, carbon steel
82 ft dia x 13.1 ft high, 574,300
527,901 gal, resin-lined,
carbon steel
Centrifugal belt drive, 5,100 51,200
gpm, 50 ft head, 125 hp,
resin-lined, carbon steel
Centrifugal belt drive, 300 10,100
gpm, 50 ft head, 10 hp,
resin-lined, carbon steel
11.5 ft dia x 14.8 ft high, 16,100
11,500 gal, open top, resin-
lined, carbon steel
6 hp, resin-lined, carbon 13,000
steel
Centrifugal belt drive, 350 32.500
gpm, 300 ft head, 60 hp,
resin-lined, carbon steel
1.743.500
Area 5—S0x-N0y. Absorption
Item
No.
Description
Total
equipment
cost. 1979 $
1. Absorbers
Tellerette packed column, 36
ft long x 36 ft wide x 33 ft
high, 316 SS
(continued)
8,816,900
45
-------
TABLE 26. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
2. Tanks, reducing
3. Agitator, reducing
tank
4. Pump, absorber
recycle
Subtotal
36 ft dia x 36 ft high,
274,100 gal, closed top,
resin-lined, carbon steel
50 hp, resin-lined, carbon
steel
Centrifugal belt drive, 15,700
gpm, 135 ft head, 1,000 hp,
resin-lined, carbon steel
623,800
203,100
353,500
9,997,300
Area 6—Crystallizing
Item
No.
Description
Total
equipment
cost, 1979 $
1. Evaporator
2. Pump, recycle first
effect evaporator
3. Pump, recycle second
effect evaporator
'4, Pump, recycle third
effect evaporator
5. Condenser, third
effect evaporator
Triple effect, 24 ft dia x 1,110,000
19.7 ft high, 66,000 gal,
316 SS
Centrifugal belt drive, 4,600 39,000
gpm, 75 ft head, 200 hp,
rubber-lined, carbon steel
Centrifugal belt drive, 4,300 39,000
gpm, 75 ft head, 200 hp,
rubber-lined, carbon steel
Centrifugal belt drive, 4,000 39,000
gpm, 100 ft head, 200 hp,
rubber-lined, carbon steel
5,300 ft2, 316 SS 472,100
(continued)
•46
-------
TABLE 26. (continued)
Item
No.
Des cription
Total
equipment
cost, 1979 $
6. Steam ejector, third
effect evaporator
7. Heat exchanger, feed- 2
stream to crystallizer
8. Crystallizer, cooling 2
9. Condenser, at
crystallizer
10. Refrigeration system
11. Steam ejector -
crystallizer
12. Pump, crystallizer
recycle
13. Separator, dithionate 13
14. Pump, overflow to 2
surge tank
15. Tank, centrate
overflow surge
16. Conveyor, dithionate 1
cake to storage
100 psig, steam consumption
1,100 Ib/hr, 1-1/2 in. syphon,
90°F suction temperature, 20
ft suction lift, resin-lined,
carbon steel
3,000 ft2, 316 SS
16.4 ft dia x 23 ft high,
37,000 gal, resin-lined,
carbon steel
25,000 ft2, air cooled,
316 SS
Cooling load, 4.87 x 106
kcal/hr
100 psig, steam consumption
2,200 Ib/hr, 1-1/2 in. syphon,
41°F suction temperature, 20
£t suction lift, resin-lined,
carbon steel
Centrifugal belt drive, 3,700
gpm, 75 ft head, 200 hp,
resin-lined, carbon steel
Screw decanter, 316 SS
Centrifugal belt drive, 1,100
gpm, 75 ft head, 40 hp, resin-
lined, carbon steel
35 ft dia x 37 ft high,
266,300 gal, open top,
resin-lined, carbon steel
Screw, 16 in. dia x 30 ft
long, 40 hp, 65 tons/hr
(continued)
5,200
132,000
73,100
1,947,000
270,000
10,400
78,000
2,225,500
18,100
129,100
10,800
47
-------
TABLE 26. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
17. Tank, dithionate
cake surge
18. Pump, overflow to
reduction reactor
19. Building
Subtotal
30 ft dia x 30 ft high,
158,600 gal, open top,
resin-lined, carbon steel
Centrifugal belt drive, 1,100
gpm, 150 ft head, 100 hp,
resin-lined, carbon steel
60 ft x 120 ft, 1 level
91,400
24,800
230,400
6,944.900
Area 7—Nitrogen Treatment
Item
No.
1. Pump, NH{S03Na)2
solution
2. Tank, NH(S03Na)2
solution surge
Agitator, NH(SQ3Na)2
solution surge tank
Pump, NH(S03Na>2
reactor feed
5. Reactor, NH(S03Na)2 1
6. Agitator, NH(S03Na)2 1
reactor
7. Separator, 1
NH(S03Na)2
Description
Total
equipment
cost. 1979 $
Centrifugal belt drive, 70 4,800
gpm, 50 ft head, 2 hp, resin-
lined, carbon steel
18 ft dia x 18 ft high, 31,800
34,300 gal, open top, resin-
lined, carbon steel
5 hp, resin-lined, carbon 23,000
steel
Centrifugal belt drive, 70 4,800
gpm, 50 ft he.ad, 2 hp,
resin-lined, carbon steel
6.6 ft dia x 14.8 ft high, 13,200
3,800 gal, 316 SS
15 hp, resin-lined, carbon 9,900
steel
Screw decanter, 316 SS 171,200
(continued)
48
-------
TABLE 26. (continued)
Item
No.
Description
Total
equipment
cost. 1979 $
8. Pump, recycle
centrate to
reducing tank
9. Conveyor, NH(S03Na>2
cake to cracker
10. Cracker, NH(SC»3Na)2
11. Conveyor, sulfate to
storage
12. Tank, K.2S04 surge
storage
13.
14.
15. Tank, H2S04 storage
Weigh feeder,
K2S04 storage
Conveyor, sulfate to
NH(S03Na)2 reactor
16. Pump,
17. Reactor, EDTA
18. Agitator-EDTA
reactor
19. Filter. EDTA
20. Pump, filtrate to
neutralization
reactor
Subtotal
1
2
Centrifugal belt drive, 70 4,800
gpm, 75 ft head, 3 hp, resin-
lined, carbon steel
Screw, 6 in. dia x 30 ft long, 4,400
7.5 hp, 3.5 tons/hr
Rotary kiln, indirect heating, 218,300
4.6 ft dia x 36 ft long, 33
kW, 310 SS
Screw, 6 in. dia x 30 ft long, 3,200
5 hp, 2.5 tons/hr
10 ft dia x 10 ft high, 59,000 9,900
gal, open top, resin-lined,
carbon steel
4,300 Ib/hr 6,400
Bucket, 6 in. x 4 in. x 4 in., 8,700
0.5 hp, 50 ft lift, 2.2 tons/hr
16 ft dia x 20 ft high, 30,100 19,800
gal, closed top, carbon steel
Centrifugal belt drive, 40 gpm, 3,600
50 ft head, 2 hp, carbon steel
7.5 ft dia x 10 ft high, 1,075 7,100
gal, open top, resin-lined,
carbon steel
3 hp, resin-lined, carbon 6,700
steel
40 ft2 filtering area 50,000
Centrifugal belt drive, 100 6,500
gpm, 50 ft head, 2.5 hp,
resin-lined, carbon steel
608,100
(continued)
49
-------
TABLE 26. (continued)
Area 8—Cracking
Item
No.
Description
Total
equipment
cost, 1979 $
1. Tank, fuel oil
storage
2. Pump, fuel oil
3. Hot blast generator
(furnace)
4. Blower, flue gas to
cracker
5. Conveyor, dithionate
cake to dryer
Weigh feeder;
dithionate surge
to dryer
Dryer
8. Conveyor, dithionate
cake to cracker
9. Cracker, dithionate
10. Conveyor, Na2SC»4 to
storage
11. Conveyor, Na2SO4 to
storage
35 ft dia x 40 ft high, 91,000
287,900 gal, closed top,
carbon steel
Centrifugal belt drive, 25 2,800
gpm, 100 ft head, 1 hp,
cast iron
35 x 106 kcal/hr, oil-fired 955,400
100,750 ft3/min, 400 hp, AP 208,500
15 in. H20
Bucket, 12 in. x 7 in. x 13,800
7-1/4 in., 10 hp, 50 ft lift,
65 tons/hr
63.3 tons/hr 15,100
Rotary kiln, direct heating, 2,389,500
9.8 ft dia x 115 ft long,
66 kW, 304 SS
Screw, 12 in. dia x 30 ft 5,800
long, 60 hp, 40.3 tons/hr
Rotary kiln, indirect heating, 3,651,500
4.6 ft dia x 36 ft long, 33
kW, 310 SS
Screw, 14 in. dia x 30 ft 5,800
long, 60 hp, 32 tons/hr
Belt, 14 in. wide x 100 ft 24,300
long, 5 hp, 20 ft lift, 32
tons/hr
(continued)
-50
-------
TABLE 26. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
12. Tank, Na2S04 surge
storage
13. Conveyor, byproduct
to storage
14. Tank, Na2S04 byproduct 1
storage
Subtotal
10 ft dia x 10 ft high, 5,900
gal, open top, resin-lined,
carbon steel
Belt, 14 in. wide x 300 ft
long, 1 hp, 5.2 tons/hr
24 ft dia x 24 ft high,
81,200 gal, open top, resin-
lined, carbon steel
11,800
68,800
68,700
7,512,800
Area 9—Gypsum Production
Item
No.
Description
Total
equipment
cost. 1979 $
1.
2.
3.
4.
5.
Heat exchanger, S02
gas
Blower, S02 No. 1
Converter, Na2S04
Agitator, Na2S04
converter
Conveyor, Na2S04 to
1
2
1
1
1
500 ft2, 316 SS
2,075 ft3/min, 10 hp, AP
15 in. H20
16.4 ft dia x 20 ft high,
31,700 gal, closed top,
resin-lined, carbon steel
30 hp, resin-lined, carbon
steel
Bucket, 8 in. x 5 in. x
42,200
3,300
75,200
34,100
13,700
premix tank
6. Weigh feeder, NaS04
to premix tank
7. Tank, premixing
5-1/2 in., 5 hp, 40 ft lift,
27 tons/hr
26.8 tons/hr
27.5 ft dia x 27.5 ft wide,
122,200 gal, open top, resin-
lined, carbon steel
(continued)
11,000
163,000
51
-------
TABLE 26. (continued)
Item
No.
Description
Total
equipment
cost. 1979 $
8. Agitator, premix 1
tank
9. Pump, Na2S04 converter 2
feed
10. Thickener, gypsum
60 hp, resin-lined, carbon 58,000
steel
Centrifugal belt drive, 360 10,AGO
gpm, 50 ft head, 10 hp, resin-
lined, carbon steel
75.5 ft dia x 16.A ft high, 380,200
A9,100 gal, resin-lined,
carbon steel
11. Pump, gypsum separator 2
feed
12. Separator, gypsum 6
13. Conveyor, gypsum to 1
storage
1A. Pump, NaHS03 tank
No. 2 feed
15. Tank, NaHS03 surge 1
No. 2
Centrifugal belt drive, 185 7,200
gpm, 50 ft head, 7.5 hp,
resin-lined, carbon steel
Screw decanter, 316 SS 578,AOO
Belt, 1A in. wide x 300 ft 76,600
long, 7.5 hp, 50 ft lift,
23.2 tons/hr
Centrifugal belt drive, 1AO 6,900
gpm, 50 ft head, 5 hp, resin-
lined, carbon steel
15 ft dia x 15 ft high, 19,800 22,500
gal, open top, resin-lined,
carbon steel
16. Pump, NaHS03 recycle 2
17. Pump, NaHS03 tank 2
No. 1 feed
18. Tank, NaHS03 surge
No. 1
19. Pump, NaHS03 feed
to S02 stripper
Subtotal
Centrifugal belt drive, 150
gpm, 50 ft head, 5 hp, resin-
lined, carbon steel
Centrifugal belt drive, 325
gpm, 50 ft head, 7.5 hp,
resin-lined, carbon steel
18 ft dia x 21 ft high, A0,000
gal, open top, resin-lined,
carbon steel
Centrifugal belt drive, 325 gpm,
100 ft head, 15 hp, resin-lined,
carbon steel
6,900
10,300
37,AOO
11,500
1.5A8.800
(continued)
52
-------
TABLE 26. (continued)
Area
1.
2.
10 — Absorbent
Item
Heat exchanger
stripper feed
Pump, SO? stri
Regeneration
No. Description
, S02 1 1.700 ft2, 316 SS
pper 2 Centrifugal belt drive,
Total
equipment
cost, 1979 $
70,600
16,500
recycle
3. Heat exchanger, S02
stripper heater
4. S02 stripper
5. Heat exchanger, S02
stripper condenser
6. Tank, NaHS03 holding
Agitator, NaHS03
holding tank
Pump, NaHSOs feed to
double decomposition
9. Heat exchanger
10. Reactor,
decomposition
11. Agitator, double
decomposition reactor
12. Pump, reactor
No. 2 feed
1
2
965 gpm, 100 ft head,
316 SS
1,400 ft2, 316 SS
16.4 ft dia x 26.2 ft high,
39,600 gal, 316 SS
700 ft2, 316 SS
32.5 ft dia x 35 ft high,
217,200 gal, open top,
resin-lined, carbon steel
50 hp, resin-lined, carbon
steel
Centrifugal belt drive, 260
gpm, 100 ft head, 20 hp,
resin-lined, carbon steel
400 ft2, 316 SS
16.4 ft dia x 19 ft high,
29,000 gal, open top,
304 SS
25 hp, resin-lined, carbon
steel
Centrifugal belt drive, 226
gpm, 50 ft head, 10 hp,
resin-lined, carbon steel
(continued)
63,800
93,600
46,800
111,200
. 49,300
11,200
40,000
166,900
60,000
8,600
53
-------
TABLE 26. (continued)
Item
No.
Description
Total
equipment
cost. 1979 $
13. Thickener, sulfite
75.4 ft dia x 16.4 ft high,
527,900 gal, resin-lined,
carbon steel
375,300
14. Pump, regenerated
solution tank feed
15. Pump, CaSOS separator
feed
16. Separator, sulfite
17. Conveyor, CaS03 cake
to tank
18. Tank, CaS03 No. 1
2 Centrifugal belt drive, 355 10,500
gpm, 50 ft head, 10 hp,
resin-lined, carbon steel
2 Centrifugal belt drive, 75 6,800
gpm, 50 ft head, 3 hp,
resin-lined, carbon steel
6 Screw decanter, 304 SS 578,400
1 Bucket, 8 in. x 5 in. x 9,700
5-1/2 in., 5 hp, 40 ft
lift, 27 tons/hr
1 25 ft dia x 20 ft high, 102,800 68,100
gal, open top, resin-lined,
carbon steel
19. Agitator, CaS03 tank
No. 1
20. Pump, CaS03 feed
No. 1
21. Tank, CaSC-3 timing
22. Agitator, CaS03
timing, tank
23. Pump, CaS03 feed
No. 2
24. Pump, No. 2 CaS03
tank feed
25 hp, resin-lined, carbon 30,000
steel
Centrifugal belt drive, 265 10,300
gpm, 50 ft head, 7.5 hp,
resin-lined, carbon steel
26 ft dia x 30 ft high, 119,100 75,500
gal, open top, resin-lined,
carbon steel
25 hp, resin-lined, carbon 30,000
steel
Centrifugal belt drive, 265 11,800
gpm, 150 ft head, 25 hp,
resin-lined, carbon steel
Centrifugal belt drive, 20 4,800
gpm, 50 ft head, 1 hp,
resin-lined, carbon steel
(continued)
54
-------
TABLE 26. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
25. Tank, CaSOs No. 2
26. Pump, CaS03 recycle
27. Tank, regenerated
solution
28. Agitator, solution
tank
29. Pump, regenerated
solution
30. Blower, S02
Subtotal
6 ft dia x 10 ft high, 2,100
gal, open top, resin-lined,
carbon steel
Centrifugal belt drive, 20
gpm, 50 ft head, 1 hp,
resin-lined, carbon steel
45 ft dia x 47 ft high,
559,100 gal, open top,
resin-lined, carbon steel
50 hp, resin-lined, carbon
steel
Centrifugal belt drive, 355
gpm, 100 ft head, 20 hp,
resin-lined, carbon steel
1,080 ft3/min, 5 hp, AP
15 in. H20
5,700
4,800
214,900
98,700
11,200
3.300
2.288.300
Area 11—Raw Material Storage
Item
No,
Description
Total
equipment
cost. 1979 S
Elevator to
storage bin
2. Tank,
storage
3. Weigh feeder,
to reducing tank
Bucket, 8 in- x 5 in. x
5-1/2 in., 10 hp, 80 ft
lift, 30 tons/hr
Cylindrical tank with cone
bottom, 37.5 ft dia x 40 ft
high, 50,254 ft3, closed top,
carbon steel
9,750 Ib/hr
(continued)
15,400
184,900
6,400
55
-------
TABLE 26. (continued)
Item
No.
Description
Total
equipment
cost. 1979 $
.4. Conveyor, Na2C03 to 1
reducing tank
5. Elevator,
to storage
6. Hopper, FeS04 storage 1
7. Weigh feeder, FeSC>4 1
to mix tank
8. Tank, Na^EDTA storage 1
9. Pump, Na4EDTA to 2
mixing tank
10. Mixing tank, Na4EDTA- 1
11. Agitator, NaAEDTA-
mixing tank
12, Pump, Fe^EDTA to reg.
solution tank
Subtotal
Total, Areas 1-11
Belt, 14 in. wide x 300 ft
long, 2,5 hp, 40 ft lift,
5 tons/hr
Bucket, 6 in. x 4 in. x
4-1/2 in., 2.5 hp, 30 ft
lift, 14 tons/hr
Square hopper with pyramidal
bottom, 10 ft long x 10 ft
wide x 10 ft high, 1,433 ft3,
carbon steel
238 Ib/hr
Cylindrical tank, 19 ft dia x
20 ft high, 42,400 gal, closed
top, resin-lined, carbon steel
Centrifugal belt drive, 60
gpm, 50 ft head, 2 hp,
resin-lined, carbon steel
5 ft dia x 5 ft high, 750 gal,
open top, resin-lined, carbon
steel
3 hp, resin-lined, carbon
steel
Centrifugal belt drive, 60
gpm, 50 ft head, 2 hp,
resin-lined, carbon steel
77,000
6,900
19,200
5,600
44,100
4,800
2,500
6,900
4.800
378,500
34,807,300
56
-------
Feed Preparation—
The feed preparation section consists of a single train of equipment
required to weigh, crush, and slurry the makeup limestone. The fresh
limestone slurry, containing 30 weight percent solids, is stored in the
limestone slurry feed tank which has a surge volume sufficient for 24 hours
normal operation.
Gas Handling—
Four ID fans, one for each train of prescrubbers and absorbers, are
provided between the reheating section and the: stack to compensate for the
pressure drop in the Asahi system. Each fan is sized to handle one-fourth
of the total flue gas volume and a pressure drop of 59 inches water in the
system.
Particulate Control—
Four trains of modified, variable-throat venturi prescrubbers and four
trains of prescrubber holding tanks are used. One train of equipment required
to concentrate, neutralize, and pump the fly ash slurry to the pond is
provided.
SOX-NOX Absorption—
Four trains of packed-bed absorbers and four reducing tanks are provided.
Six recycle pumps (four operating and two spares) are provided to recirculate
the scrubbing solution from the reducing tanks to the absorbers.
Crystallization Section—
Two triple-effect evaporator systems to concentrate the absorber purge
solution and two cooling crystallizers are provided.
N32S04 and Na2S206 crystals are separated in 13 screw decanters
operating in parallel. The solids storage tank has an 8-hour capacity and
the liquid storage tank has a 4-hour capacity.
Nitrogen Treatment Section—
A single train of processing equipment is used to separate and process
NH(S03Na)2.
Cracking Section—
This section contains a single train of equipment to first dry and then
thermally decompose the sodium salts from the crystallization section. An
oil-fired furnace is provided for heating. A byproduct Na2SO^ storage tank
with a 7-day storage capacity is provided.
Gypsum Production—
Since makeup Na2C03 is relatively expensive, this section has been
included to recover the sodium and, indirectly, to convert the sulfate salt
to gypsum. In addition to the equipment to produce gypsum, the equipment
needed to pass the byproduct S02 from the crackers through the gypsum reactor
to form additional NaHS03 has been included.
57
-------
Absorbent Regeneration—
Since ^2803, rather than the NaHSC>3 formed in the previous gypsum
production section, is the active compound in reducing the absorbed NOX, the
single train of equipment in this section is provided to convert the bisulfite
to reusable sulfite in two stages. Initially the solution is heated to
strip any excess S02 and is then neutralized with limestone slurry to form
both Na2SC>3 and CaSO-j. The CaSC>3 is separated, reslurried, and sent back
to the gypsum production section. The Na2SC>3 is recycled to the reducing
reactor.
Raw Material Storage—
This section includes all the processing equipment needed to receive,
store, and retrieve for later use all of the various raw materials (except
limestone) needed in the process. Storage capacity is designed for 30
days of normal usage.
IHI SO^-NtX. PROCESS - WET, OXIDATION-ABSORPTION-REDUCTION
A. A
Process Description
The IHI process oxidizes NO to N02 by injecting ozone into the flue gas
stream (11). This N02 is then absorbed into a calcium-based slurry and reacts
with the absorbed S02 to form a mixture of nitrate salts, molecular nitrogen,
and complex nitrogen-sulfur compounds. The complex nitrogen-sulfur compounds
are decomposed into molecular nitrogen and calcium sulfate (CaS04). Most of the
S02 in the flue gas absorbed into solution as sulfite ions are oxidized to
sulfates in an oxidizing tower and removed in a centrifuge as "byproduct gypsum.
The overall system is very similar to the other wet, oxidation-absorption-
reduction processes, in that it consists of six major sections: (1) prescrub-
bing, (2) gas-phase oxidation, (3) absorption, (4) oxidation of sulfites and
reduction of absorbed NOX, (5) byproduct gypsum recovery, and (6) decomposi-
tion of complex nitrogen-sulfur compounds.
The. flue gas from the air heater passes through a prescrubber which
removes most of the particulates and essentially all of the chlorides. The
flue gas is cooled from 300°F to 127°F and humidified by the evaporation of
water from the scrubbing solution. Part of the scrubber liquid is purged
to a thickener and then to a filter. The filtrate is recycled to the thick-
ener. Most of the thickener overflow is recycled to the prescrubber after
a small purge stream has been removed for chloride control. This purge
stream passes through a secondary thickener to remove any remaining fly ash
and, after being used as wash water for the byproduct gypsum cake, is pumped
to the bottom of the absorber.
As the cooled flue gas passes from the prescrubber to the absorber, ozone
in a weak ozone-air mixture is injected to selectively oxidize the NO to N02
by the following reaction:
N°(g) + °3 * ""'(g) + °2(g)
58
-------
Since an 03:NOX mol ratio of only 0.85:1.0 is used for 90% NOX removal, the
ozone injection system must be properly designed to give good mixing characte1--
istics and to prevent the undesirable formation of dinitrogen pentoxid^ (N20s) :
2N02 + 03 + N205 + 02 (2)
(g) (g) (g)
Following oxidation, the flue gas is contacted with a limestone slurry
in a specially designed countercurrent-f low spray tower. This slurry has a
pH of 4-6 and contains low concentrations of copper chloride (CuCl2) and
sodium chloride (NaCl) as catalysts to aid in the absorption and reduction
of N02. The S02 is absorbed from the flue gas and undergoes the following
liquid-phase reactions:
802 (8) * S°2(aq) (3)
S°2(aq) + CaC°3(s) + 1/2H2° * CaS03-l/2H20(8) + C02(g) (4)
s°2(aq) + CaS03-l/2H20(s) + 1/2 H20 •*• Ca(HS03)2 (5)
The N02 is absorbed and undergoes the following liquid-phase reactions:
2N02(g) - N20A(aq) (6)
N2°4(aq) + 4CaS03-l/2H20(g) + 6H20 •* N2(g) + 4CaS04'2H20(s) (7)
N2°4(aq) + Ca(HS°3)2(aq) + 5CaS03.l/2H20(g) + 17/2H20 ->
Ca(NH2S03)2(a{j) + 5CaS04-2H20(s) (8)
2N204(aq) + 4CaS03-l/2H20(s) +
Ca(N02)2(aq) + Ca(N03)2(aq) + 2Ca(HS03)2(aq) (9)
In addition to these primary reactions, the other secondary reactions
occur in the absorbing solution. Oxygen absorbed from the flue gas oxidizes
sulf ite ion .by the following reaction:
CaS03-l/2H20(s) + l/202(aq) + 3/2H20 •* CaS04-2H20(s) (10)
The NOx, including some N2(>5, can also undergo the following reactions:
+ 3CaS03.l/2H20(g) + 3/2H20 •* N20(g) + 3CaS04-2H20(s) (11)
N2°5(aq) + 2CaS03.l/2H20(g) - Ca(N03)2(aq) + Ca(HS03)2(aq) (12)
59
-------
Thus the S02, absorbed to form the sulfite ion, is used to reduce the N02
to molecular nitrogen and complex nitrogen-sulfur compounds. Approximately
80% of the absorbed NOX is converted to calcium sulfamates [Ca(NH2S03) 2! and
most of the remainder to molecular nitrogen. According to the process developer
nitrate salts represent only a small amo'unt (<10%) of the total absorbed NOX.
After passing through a mist eliminator, the cleaned flue gas is reheated
to provide plume buoyany and is then exhausted through the stack.
From the absorber, the absorbent slurry flows into a holding tank where
fresh catalyst and makeup limestone slurry are added. Most of the absorbent
slurry is recycled to the absorber. A small (1-2% of the circulating slurry)
purge stream is acidified with a small amount of H2S04- The acidified slurry
is pumped to an oxidizing tower where the CaSC>3 in the purge stream is con-
verted to gypsum as it passes countercurrently through an airstream. The
gypsum slurry is pumped to the gypsum thickener for initial separation.
The thickener bottoms are centrifuged to remove byproduct gypsum and the
centrate is recycled to the gypsum thickener.
The overflow from the gypsum thickener, which contains complex nitrogen-
sulfur compounds and nitrate salts is neutralized with powdered limestone to
precipitate the dissolved catalyst and pumped to a thickener. The bottoms
from the thickener, mostly excess limestone and catalyst, are recycled to the
absorber. The thickener overflow is punped to a second neutralization reactor
where additional powdered limestone is added to prevent the formation qf NH3.
The effluent from the reactor is passed through a double-effect evaporator
system to form a mixed salt cake. This cake is passed through a rotary kiln
operating at 147-165°F where the sulfamates and the nitrate salts are decom-
posed to molecular nitrogen and CaSC>4. The decomposition of the Ca(NH2S03>2
salts is given by the following reaction:
2Ca(NH2S03)2(aq) + 4CaC03(aq) + 2Ca(HS04) 2(aq) + 302(aq) + 8H20 •>
2N2(g) + 8CaS04-2H20(s) + ^02(g) (13)
The off-gas from the decomposer, which may contain some S02 in addition to
the nitrogen and carbon dioxide (C02) , is recycled to the flue gas ducts.
Analysis of Processing Subsections
The flow diagram and material balance for the IHI process in the base
case application are shown in Figure 3 and Table 27 respectively. The IHI
SOX-NOX system is composed of 10 separate processing sections. The detailed
equipment list for the IHI process listed by processing section is shown in
Table 28. Each of these subsections is described in more detail in the
following subheadings .
Material Handling —
This and the following area, fe.ed preparation, comprise the limestone
preparation section. The material-handling section includes all equipment
60
-------
H(S04
»
T*
ffi
k
i
SO
IS
LS
Am
^
11 /
1
» ri
-^ c> " •.
CENTRIFUGE
NaCI
SILO
AIR MAKEUP
CiiClj-
9 1
CE
SYPSUW
TO DISPOSAL
NEUTRALIZATION NCUTRALI2ATION
| REACTOR I I REACTOR I
(CATALTST ffeeoVtMY I (NHy SUFMEBSIOM I
EVAPORATOR UNIT
STEAM
3}
S I
-•
HOPPERS. FEEDERS. A CONVETOHS
Figure 3. IHI SOX-NOX process. Flowsheet.
-------
TABLE 27. IHX SOX-"0X PROCESS
MATERIAL BALANCE
Stream No.
Description
\
2
')
It
',
6
1
H-
9
1°
Total stream, Ib/hr
Sft3/oin (60°F)
Temperature , UP
Pressure, psig
Gpra
Specific gravity
pH
Undlssolved solids, %
1
Coal to boiler
428,600
1
Combustion air
to air heater
4.546,200
1,005,000
80
3
Combustion air
to boiler
4,101,800
906,700
535
4
Gas to
economizer
4.514.100
958.000
890
5
Gas to
air heater
4.516.100
958.000
705
Stream No.
Description
1
2
i
4
"S
h
7
a
9
l(?
Total stream, Ib/hr
Sft3/min (60'F)
Temperature. °F
Pressure. psis
Gum
Specific gravity
pH
Undissolved solids, %
6
Gas to
prescrubber
4.960.400
1.056.000
300
7
Makeup H20
287,900
575.3
8
Recycle solution
to the top of
the prescrubber
N7A
127
1.02
1.5
2.0
9
Prescrubber
purge to ash
removal
2,860,000
127
1.02
1.5
2.0
10
Fly ash
thickener
underflow
613,800
1,160
1.06
8.9
Stream No.
Description
1
i
4
5
h
7
H
y
10
Total stream. Ib/hr
Sft3/min (60°F5
Temoerature. °F
Pressure, osie
Com
Snecific aravitv
pH
Undissolved solids. %
11
Fly ash slurry
feed to
filter
616.000
127
1.161
1.06
8.9
12
Fly ash to
disposal
68.300
127
80
13
?iltrate recycle
to fly ash
thickener
547,700
127
0.0
14
?Iy ash
thickener
overflow
2,794,000
127
5,556
1.00
0.095
IS
Fly ash
thickener
overflow
to prescrubber
2,728,100
Stream No.
2
i
h
/
H
9
10
Description
Total stream. Ib/hr
SftJ/ttin (SO^F)
' Temperature. °F
Pressure. psiE
CpTO
Specific gravity
pH
Undissolved solids, %
16
Fly ash
thickener
overflow to
secondary
thickener
66,000
n.rws
1 17
Secondary
thickener
underflow to
holding tank
2,200
1.02
4.1
1 I3 1
Secondary
thickener over-
flow recycled to
absorber (via
;ypsum centrifuge
65,800
19
Flue gas from
i prescrubber
5,115,700
1.070.200
127
20
Clean air feed
to ozone
401,800
88,359
(continued)
62
-------
TABLE 27. (continued)
Description
1
1
h
H
9
10
Total stream, Ib/hr
Sft3/min (60°F)
Temperature, °F
Pressure, psig
Gprn
Specific gravity
pH
Undlssolved solids, %
21
Air-ozone
mixture from
generators
403,800
22
Flue gas-ozone
mixture
5,519,632
1,155,805
127
23
Flue gas to
absorber
5,554,396
1,163,085
127
Flue gas to
reheator
5.567.314
1.164.625
127
-i
Flue gas to
stack
5. 567. 314
1.165.790
175
1
j
1
^
ii
H
>i
12.
Total stream, Ib/hr
Sfti/min (60°F)
Temperature, °F
Pressure, pslg
Gpn
Specific gravity
pH
Undissolved solids, %
26
Makeup NaOH
620
27
Makeun CuCl2
22
28
Makeup SaCl
330
29
Makeup H?0
53.600
60
103
30
Mixing tank
to absorber
54,600
Description
1
j
i
.,
i
h
-
H
1
10
Total stream, Ib/hr
Sfti/mln (60°F)
Temperature, °F
Pressure, psig
Gpm
Specific gravity
PH
Undiseolved solids, Z
31
Absorber holding
75. 372. 000
127
1.17
7.93
Recycled
absorbent to
the top of the
absorber
74,800,000
127
1.17
7.93
Absorbent
solution to
regeneration
section
572,000
127
1.17
7.93
Makeup H2S04
220
60
M*
Forward feed
to oxidation
tower
572.220
1.17
Stream No.
1
*
t
4
-i
h
7
H
q
Total stream, Ib/hr
Sft-Vmln <60°n
Temperature, °F
Pressure, psig
Gpm
Specific gravity
pH
Undig solved solids, %
36
Air feed to
oxidation
tower
15.510
3.394
60
Forward feed
to gypsum
thickener
574.400
127
n 1.18
3.5
8.7
Gypsum
thickener
overflow to
holding tank
517.200
Gypsum
thickener
overflow
recycled
to absorber •
246.400
40
Gypsum thickener
overflow to
Nj-treatment and
limestone
preparation
270, aoo
{continued)
63
-------
TABLE 27. (continued)
Description
|
2
\
•t
h
7
H
M
|0
Total stream, Ib/hr
Sft3/min (60°F)
Temperature, °F
Pressure, psig
Ggra
Specific gravity
PH
Undissolved solids, Z
41
Gypsum
thickener
underflow
171.600
1.33
30
42
Byproduct
gypsum to
disposal
57.200
84.3
43
Centrate
recycled to
gypsum
thickener
114.400
O.Q
44
Overflow to
limestone
slurry feed
tank and wet
ball mill
184,000
45
Gypsum
thickener
overflow to No.l
neutralization
reactor
86,000
Stream No.
Description
1
2
)
4
5
h
7
H
4
10
Total stream, Ib/hr
Sft3/min (60°F)
Temperature, °F
Pressure, psig
Gpm
Specific gravity
pH
Undissolved solids, %
46
Effluent from
neutralization
reactor
86,100
47
Thickener
underflow
recycled to
absorber
4,400
48
Thickener
overflow to
neutralization
reactor No. 2
81,700
49
Evaporator feed
84,900
3.8
50
Thermal
decomposer feed
13.500
212
Stream No.
Description
1
•j
i
4
5
h
'1
H
9
10
Total stream. Ib/hr
Sft3/min (60°F)
Temperature. °F
Pressure, usie
Gem
Specific gravity
cH
Undissolved solids, %
51
Mixed salts
to disposal
10,200
1.562
52
Fuel oil to
furnace
1.310
53
Combustion air
to furnace
27.202
5.752
60
54
Combustion
flue gas
to decomposer
28.510
5.971
1,562
55
Off -gas thermal
decomposer
31.900
6.885
1.562
Stream Mo.
Description
1
J
>
4
'•>
h
7
H
9
10
Total stream. Ib/hr
Sft3/min (60°F)
Temoerature. °F
Pressure. nsiE
Gpin
Specific gravity
DH
Undissolved solids, 7,
56
Not used
57
Makeup powdered
limestone to
neutralization
reactor No. 1
140
58
Makeup powdered
limestone to
neutralization
reactor No, 2
3,230
59
Makeup
limestone
from pile
33,700
60
Gypsum
thickener
overflow to wet
ball mill
21.700
(continued)
64
-------
TABLE 27. (continued)
Stream No.
Description
|
?
i
ri
h
]
K
t)
12.
Total stream. Ib/hr
SftJ/min (60°)
Pressure, pslg
Gpm
Specific gravity
pH
Undlssolved solids, "/.
61
Wee ball mill
product to
product tank
54f400
1.73 '
60
62
Gypsum
thickener
overflow to
limestone
slurry tank
163,100
63
Fresh makeup
limestone
slurry to
absorber
2:7,500
1.23
15
h
7
S
T
10
10
65
-------
TABLE 28. IHI SOX-NOX PROCESS
BASE CASE EQUIPMENT LIST
DESCRIPTION AND COST
Area
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
13.
14.
1 — Material Handling
Item
Car shaker and hoist
Car puller
Hopper, unloading
Feeder, unloading
vibrating
Conveyor, unloading
belt
Conveyor, unloading
incline belt
Unloading pit dust
collector
Pump, unloading pit
sump
Conveyor, storage
belt
Tripper, storage
conveyor
Mobile equipment
Hopper, reclaim
Feeder, reclaim
vibrating
Conveyor, reclaim
No.
1
1
1
1
1
1
1
1
1
1
1
1
1
1
Description
20 hp shaker, 7.5 hp hoist
25 hp puller, 5 hp return
16 ft dia x 10 ft straight
side height, carbon steel
3.5 hp
20 ft horizontal, 5 hp
310 ft, 50 hp
Polypropylene bag type,
2,200 ft-Vmin, 7.5 hp
60 gpm, 70 ft head, 5 hp
1,000 ft, 2 hp
30 ft/min, 1 hp
Scraper tractor
7 ft wide x 4.25 ft high,
2 ft wide bottom, carbon
steel
3.5 hp
100 ft, 2 hp
Total
equipment
cost, 1979 $
31,000
52,000
12,800
14,400
17,700
89,200
19,200
4,200
45,200
16,400
137,800
1,500
14,400
32,200
belt
(continued)
66
-------
TABLE 28. (continued)
15.
16.
17.
18.
19.
Area
1.
2.
3.
A.
5.
6.
7.
Item
Conveyor, reclaim
incline belt
Dust collector,
reclaim pit
Pump, reclaimer
pit sump
Elevator, reclaim
bucket
Bin, feed
Subtotal
2 — Feed Preparation
Item
Feeder, weigh bin
Gyratory crushers
Ball mill dust
collectors
Ball mill
Tank, ball mill
product
Agitator, mills
product tank
Pump, ball mill
product tank
No.
1
1
1
1
2
No.
2
2
2
2
2
2
3
Description
193 ft, 40 hp
Polypropylene bag type
60 gpm, 70 ft head, 5 hp
90 ft high, 75 hp
13 ft dia x 21 ft straight
side height, covered, carbon
steel
Description
14 ft pulley centers, 2 hp
75 hp
Polypropylene bag type, 2,200
ft3/min, 7.5 hp
9.3 tons/hr, 125 hp
10 ft dia x 10 ft high, 5,500
gal, open top, flakeglass-
lined, carbon steel
10 hp
Centrifugal belt drive, 35
gpm, 50 ft head, 1.5 hp,
resin-lined, carbon steel
%
(continued)
67
Total
equipment
cost, 1979 $
53,800
19,200
4,200
86,400
33,200
684,800
Total
equipment
cost, 1979 ^
39,600
119,500
38,400
320,100
26,800
17,500
7,100
-------
TABLE 28. (continued)
Item
No.
Description
Total
equipment
cost. 1979 $
8. Tank, limestone
slurry feed
9. Pump, limestone
slurry feed
10. Agitator, limestone
slurry tank
Subtotal
31 ft dia x 32 ft high,
180,700 gal, open top,
resin-lined, carbon steel
Centrifugal belt drive, 355
gpm, 100 ft head, 20 hp,
resin-lined, carbon steel
60 hp, resin-lined, carbon
steel
100,900
11,500
58,700
740,100
Area 3—Gas Handling
Item
1. Blower, flue gas
Subtotal
No. Description
2 772,000 aft3/min, 4,000 hp,
AP 21 in. H20, 316 SS
Total
equipment
cost, 1979 $
1,596,300
1,596,300
Area 4—Particulate Control
Item
1. Prescrubber
No.
1
Description
Venturi with sp
ray, 64.9 ft
Total
equipment
cost, 1979 $
4,788,600
2. Pump, prescrubber
recycle
dia x 105 ft high, resin-
lined, carbon steel
Centrifugal belt drive,
13,700 gpm, 75 ft head, 500
hp, resin-lined, carbon
steel
(continued)
187,900
68
-------
TABLE 28. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
3. Thickener, primary
fly ash
4. Pump, primary
thickener underflow
5. Tank, filter feed
6. Agitator, holding
tank
7. Pump, filter feed
8. Ash filter
9. Pump, filtrate
recycle
10. Pump, thickener
overflow to
holding tank
11. Tank, primary
thickener over-
flow holding
12. Agitator, holding
tank
13. Pump, secondary
thickener feed
98.4 ft dia x 11.6 ft high,
88,200 gal, resin-lined,
carbon steel
Centrifugal belt drive, 1,160
gpm, 75 ft head, 40 hp, resin-
lined, carbon steel
35 ft dia x 40 ft high,
287,900 gal, open top,
resin-lined, carbon steel
50 hp, resin-lined, carbon
steel
Centrifugal belt drive, 1,165
gpm, 50 ft head, 30 hp, resin-
lined, carbon steel
Vacuum, 485 ft'
area
filtering
Centrifugal belt drive, 1,090
gpm, 50 ft head, 25 hp,
resin-lined, carbon steel
Centrifugal belt drive, 5,600
gpm, 75 ft head, 200 hp,
resin-lined, carbon steel
48 ft dia x 50 ft high,
676,800 gal, open top,
resin-lined, carbon steel
35 hp, resin-lined, carbon
steel
Centrifugal belt drive, 5,600
gpm, 75 ft head, 200 hp,
resin-lined, carbon steel
(continued)
650,500
19,300
60,200
51,100
16,300
819,000
15,800
68,900
243,500
39,900
68,900
69
-------
TABLE 28. (continued)
Item
No,
Description
Total
equipment
cost, 1979 $
14. Thickener, secondary 1
fly ash
15. Pump, secondary 2
thickener underflow
16. Pump, secondary
thickener overflow
17. Conveyor, fly ash to
reslurrying tank
18. Tank reslurrying
19. Agitator, reslurrying 1
tank
20. Pump, fly ash slurry 4
21. Pump, pond return
Subtotal
21.3 ft dia x 9.88 ft high,
26,200 gal, resin-lined,
carbon steel
Centrifugal belt drive, 5 gpm,
75 ft head, 0.25 hp, resin-
lined, carbon steel
Centrifugal belt drive, 130
gpm, 50 ft head, 3 hp motor,
resin-lined, carbon steel
Belt, 14 in. wide x 100 ft long,
7.5 hp, 30 ft lift, 34.1 tons/hr
22 ft dia x 24 ft high, 68,200
gal, open top, resin-lined,
carbon steel
20 hp, resin-lined, carbon
steel
Centrifugal belt drive, 665
gpm, 100 ft head, 40 hp,
resin-lined, carbon steel
Centrifugal belt drive, 600
gpm, 100 ft head, 25 hp,
resin-lined, carbon steel
100,500
4,600
6,700
25,500
28,000
27,700
26,800
19,700
7.269,400^
Area 5—SOX-NOX Absorption
Item
No.
Description
Total
equipment
cost. 1979 1
1. Absorber
Spray tower, 53.5 ft dia x
139.4 ft high, resin-lined,
carbon steel
(continued)
4,268,300
70
-------
TABLE 28. (continued)
Item
2. Pump, absorber
recycle
Subtotal
Area 6 — Reheat
Item
1. Reheater
2. Soot blowers
Subtotal
Area 7 — Gypsum Production
Item
Total
equipment
No. Description cost, 1979 $
9 Centrifugal belt drive, 21,500 966^100
gpm, 75 ft head, 800 hp,
resin-lined, carbon steel
5,234,400
Total
equipment
No. Description cost, 1979 $
1 Indirect steam, 20,000 ft2, 636,200
316 SS
20 Air, fixed 130,000
766,200
Section
Total
equipment
No. Description cost, 1979 $
1. Tank, H2SOA storage 1 12 ft dia x 15 ft high, 12,700 11,100
2. Pump, H2S04 feed
Tank, slurry-acid
mixing
Agitator, slurry-acid
mix tank
gal, closed top, carbon steel
Centrifugal belt drive, 0,25 4,500
gpra, 50 ft head, 0.25 hp,
carbon steel
22 ft dia x 24 ft high, 68,100 53,200
gal, open top, resin-lined,
carbon steel
25 hp, resin-lined, carbon 31,000
steel
(continued)
71
-------
TABLE 28. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
5. Pump, oxidation
tower feed
6. Blower, air feed to
oxidation tower
7. Oxidation tower
8. Pump, gypsum
thickener feed
9. Thickener, gypsum
10. Pump, gypsum
thickener underflow
11. Separator, gypsum
12. Pump, centrate
recycle
13. Pump, wash solution
recycle
14. Conveyor, gypsum
to outdoor storage
15. Pximp, gypsum
thickener overflow
Subtotal
21
2
Centrifugal belt drive, 980
gpm, 75 ft head, 40 hp,
resin-lined, carbon steel
3,400 aft3/min, 12.5 hp, AP
15 in. H^O, carbon steel
Self-supporting, vertical
tower with smoke atomizers,
19.7 ft dia x 42.6 ft high,
flakeglass-lined, carbon steel
Centrifugal belt drive, 975
gpm, 50 ft head, 25 hp,
resin-lined, carbon steel
50.8 ft dia x 10.4 ft high,
157,100 gal, resin-lined,
carbon steel
Centrifugal belt drive, 260
gpm, 50 ft head, 7.5 hp,
resin-lined, carbon steel
Centrifuge
Centrifugal belt drive, 205
gpm, 50 ft head, 5 hp,
resin-lined, carbon steel
Centrifugal belt drive, 130
gpm, 75 ft head, 4 hp,
resin-lined, carbon steel
Belt, 14 in. wide x 300 ft long,
10 hp, 50 ft lift, 28.6 tons/hr
Centrifugal belt drive, 925
gpm, 100 ft head, 50 hp,
resin-lined, carbon steel
14,500
5,300
67,100
13,000
274,100
10,300
1,127,700
7,000
7,000
11,400
15,300
1,652,500^
(continued)
72
-------
TABLE 28. (continued)
Area 8—Nitrate Treatment
Item
No,
Description
Total
equipment
cost, 1979 $
1. Reactor, neutraliza-
tion catalyst
recovery
2. Agitator, neutraliza-
tion reactor No. 1
3. Pump, effluent feed
to thickener
4. Thickener, catalyst
recovery
Pump, thickener under- 2
flow recycle
Pump, thickener
overflow
7. Reactor, neutraliza-
tion reactor NH3
suppression
8. Agitator, neutraliza-
tion reactor No. 2
9. Pump, effluent feed
to evaporator No. 1
10. Evaporator
11. Conveyor, evaporator 1
to decomposer
22,5 ft dia x 25 ft high, 58,900
23,700 gal, open top,
resin-lined, carbon steel
20 hp, resin-lined, carbon 27,700
steel
Centrifugal belt drive, 155 7,000
gpm, 50 ft head, 4. hp,
resin-lined, carbon steel
21.3 ft dia x 34.7 ft high, 167,600
92,800 gal, resin-lined,
carbon steel
Centrifugal belt drive, 10 gpm, 4,600
75 ft head, 0.5 hp, resin-
lined, carbon steel
Centrifugal belt drive, 150 7,000
gpm, 50 ft head, 4 hp,
resin-lined, carbon steel
22 ft dia x 25 ft high, 22,600 55,100
gal, open top, resin-lined,
carbon steel
20 hp, resin-lined, carbon 27,700
steel
Centrifugal belt drive, 150 7,000
gpm, 50 ft head", 4 hp, resin-
litied, carbon steel
Double effect, 7.5 ft dia x 1,985,200
8 ft high, 2,600 gal, 316 SS
Belt, 14 in. wide x 100 ft long, 25,100
1 hp, 30 ft lift, 7.0 tons/hr
(continued)
73
-------
TABLE 28. (continued)
Item
No.
Description
Total
equipment
cost. 1979 $
12. Thermal decomposer
13. Blower, decomposer
off-gas
14. Conveyor, mixed salts
to temporary storage
15. Conveyor, mixed salts
to disposal
16. Tank, fuel oil
storage
17. Pump, fuel oil
18. Oil-fired furnace
Subtotal
Rotary kiln, direct heat, 460,000
5.9 ft dia x 82 ft long,
316 SS
26,800 aft3/min, 100 hp, AP 17,600
14 in. H20, 316 SS
Screw, 14 in. x 40 ft long, 5,100
15 hp, 10 tons/hr
Belt, 14 in. wide x 300 ft 79,300
long, 1.5 hp, 50 ft lift,
5.1 tons/hr
27.5 ft dia x 28 ft high, 93,900
124,200 gal, closed top,
carbon steel
Centrifugal belt drive, 3 4,600
gpm, 100 ft head, 0.25 hp
27 MBtu/hr heat load 167,000
3,200,400
Area 9—Raw Material Storage
Item
No.
Description
Total
equipment
cost. 1979 $_
1. Tank, NaOH storage
2. Pump, NaOH feed
3. Tank, mixing
15 ft dia x 20 ft high,
26,400 gal, closed top,
resin-lined, carbon steel
Centrifugal belt drive, 0.6
gpm, 50 ft head, 0.26 hp,
resin-lined, carbon steel
20 ft dia x 22 ft high,
51,700 gal, open top,
resin-lined, carbon steel
(continued)
19,100
4,600
44,600
74
-------
TABLE 28. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
4.
5.
Agitator, mixing
tank
Pump, additive
1
2
15 hp, resin-lined, carbon
steel
Centrifugal belt drive, 110
21,600
107,400
solution feed
6. Conveyor, feed to
limestone silo
7. Silo, powdered
limestone
8. Weigh feeder, 1
powdered limestone
9. Conveyor, limestone to 1
neutralization reactor
No. 2
Subtotal ____^_
gpm, 50 ft head, 2.5 hp,
resin-lined, carbon steel
Bucket, 6 in. x 4 in. x 4-1/4
in., 3 hp, 75 ft lift, 2.0
tons/hr
Conical bottom, 20 ft dia x
40 ft high, reinforced
concrete
3,300 Ib/hr
Belt, 14 in. wide x 100 ft
long, 0.5 hp, 25 ft lift,
1.6 tons/hr
10,300
66,000
6,300
24.800
304.700
Area 10—Ozone Generating System
Item
1. Ozone generation
package
Subtotal
Total, Areas 1-10
No.
Description
Total
equipment
cost/1979 $
Includes all equipment except 16,984.000
instrumentation and controls
76.984.000
98,432,800
75
-------
needed to receive and unload limestone from rail or truck transport and the
equipment needed to supply limestone to the weigh feeders. Only one train of
processing equipment is used.
Feed Preparation—
The feed preparation area consists of one train of equipment to weigh,
crush, and slurry the makeup limestone. The fresh limestone slurry, con-
taining 15 weight percent solids, is stored in the limestone slurry feed tank
with a surge volume sufficient for 8 hours of normal operation.
Gas Handling—
Two ID fans are provided between the reheating systems and the stack to
overcome the pressure drop through the IHI SOx~NOx process. Each fan has
been sized to handle one-half of the total flue gas volume at a pressure drop
of 21 inches water.
Particulate Control—
A single venturi-spray tower is used. The venturi is designed with
sufficient liquid surge volume to make a prescrubber holding tank unnecessary.
A single train fly ash slurry thickener and filter system is used.
SOX-NOX Absorption—
Simultaneous SOx-NOx absorption is carried out in a single spray tower,
designed with sufficient liquid retention time in the base of the tower to
make a separate absorbent holding tank unnecessary.
Reheat—
Flue gas reheat to 175°F is provided by a single reheater using an
indirect steam system.
Gypsum Production—
During normal operation most of the S02 is converted to CaS03 and only
some is oxidized to gypsum. Since this mixture will not form a solid material
without further treatment, the purge from the absorber is air-oxidized to
convert the sulfite-sulfate mixture to sulfate.
Nitrogen Treatment Section—
Since significant amounts of the absorbed NOx are converted to both
nitrate salts and complex nitrogen-sulfur compounds, a purge stream con-
taining these mixed salts is neutralized to remove the dissolved catalyst,
evaporated, and the resulting mixed salts thermally decomposed. The elevated
temperatures for decomposing these mixed salts are obtained by using the flue
gas from an oil-fired furnace. This purge stream is relatively small and the
equipment is designed so that only one train of equipment is required.
Raw Material Storage—
This section includes all the processing equipment needed to receive,
store, and retrieve for later use all of the various raw materials (except
limestone) needed in the process. Storage capacity is designed for 30 days
of normal usage.
76
-------
Oxone Generation Section —
The ozone system is designed as a package unit to include all of the
process equipment needed to generate ozone from air (13).
MORETANA CALCIUM SOX-NOX PROCESS - WET, OXIDATION-ABSORPTION-REDUCTION
Process Description
The Moretana Calcium process is a wet, oxidation-absorption-reduction
system using C102 for the gas-phase oxidation of NO to NQ2 (1). Although
the use of C102 results in the formation of undesirable nitrate and chloride
salts in the scrubbing solution, it has the advantages of costing less to
produce and of requiring a lower oxidant to NOX mol ratio. The overall
process consists of five basic steps: (1) prescrubbing, (2) gas-phase oxida-
tion, (3) simultaneous S0x-N0x absorption, (4) byproduct sludge removal, and
(5) byproduct nitrate removal.
The flue gas from the air heater initially passes through an ESP which
removes 94% of the fly ash. Further fly ash removal as well as chloride
removal and humidif ication is achieved by passing the flue gas through a
prescrubber countercurrent to an aqueous scrubbing solution. The liquid
effluent from the prescrubber is recirculated after makeup water has been
added. A small purge stream of scrubbing liquid is pumped to a pH-controlling
tank where a slaked lime slurry is used to neutralize both the prescrubber
purge and a waste stream from the C102 plant. The resulting effluent from
the pH tank is sent to the fly ash pond.
The flue gas, after passing through a mist eliminator, is injected with
a C102-air mixture. The C102 is generated onsite and is injected at a rate
of 10% in excess of the stoichiometric requirement. With the proper design
of the injection system, the C102 selectively oxidizes the NO by the
following reaction:
2N°(g) + Cl°2(g) + H2°(g) - N°2<8) + ^(S) + HC1(8) (1)
This gas-phase reaction occurs very rapidly and completely oxidizes the NO
to N02 and
The oxidized flue gas is then passed countercurrently to a limestone
slurry in a Moretana plate tower absorber. The limestone slurry has a pH
of approximately 5.5. It is a mixture of calcium salts, mainly CaS03, and
contains a catalyst to aid in the reduction of NOx. As the slurry passes
through the absorber, the S02 is stripped from the flue gas and the following
reactions occur in the scrubbing solution:
S02(g) - S02(aq) (2)
CaC03(g) + S02(aq) - CaS03(aq) + C02(g) (3)
CaS°3(aq) + S°2(aq) + H2° * Ca(HS03>2(aq)
77
-------
At the same time the gases formed from the oxidation of NO are also
absorbed from the flue gas stream and undergo the following hypothesized
reactions:
2N02(g) + N204(g) (5)
(6)
q) + 2CaS°3(aq) + 2CaC03(s) + 2H2° Cat*
2Ca(N02)2(aq) + 2CaSOv2H20(s) + 2C02(g) (7)
Ca(HS03)2(aq) + Ca(N02)2(aq) + H20 cat^yst N2(g) + 2CaS04 -2H20(s) (8)
HN03(g) -> HN03(aq) (9)
CaC03(g) + 2HN03(aq) -> Ca(N03)2(aq) + H20 + C02(g) (10)
HC1, . -»• HC1, x (11)
(g) (aq)
2HC1(aq) + CaC°3(s) * CaC12(aq) + C02(g) + H2° <12)
Secondary reactions also occur in the absorber; for example, dxygen
from the flue gas is absorbed into the scrubbing solution and causes the
oxidation of CaS03 to gypsum:
CaS°3(aq) + 1/2°2(aq) + 2H2° * CaS04 ' 2H2°(S) (13)
However, this oxidation is not as serious a problem in the calcium system as
in the sodium systems because CaS03 is relatively insoluble in the scrubbing
solution. Hence, most of the CaS03 is not present in the solution in a
readily oxidizable form. After passing through a mist eliminator, the
cleaned flue gas is reheated and vented to the stack.
The liquid effluent from the absorber is recirculated through the
absorber after fresh makeup limestone slurry has been added. A small purge
stream of the slurry is pumped to a centrifuge where the relatively insoluble
CaS03 and CaS04 are separated from the soluble calcium chloride (CaCl2)
and Ca(N03)2- The solids are reslurried and pumped to a sludge pond. The
centrate containing the soluble calcium salts is separated into two streams,
the largest of which is pumped to the limestone preparation section to trans-
port makeup limestone and makeup catalyst to the absorber holding tank. The
second stream is purged to a byproduct nitrate treatment section.
78
-------
The mixed solution of calcium salts is treated by reacting the solution
with (NH4)2S04 to produce a liquid fertilizer by the following reactions:
CaCl2(aq) + (NH4)2S04(aq) + 2H20 + 2NH4C1 (aq) + CaS04 -2H20(s) (14)
Ca(N03)2(aq) + (NH4)2S04(aq) + 2H20 -> 2NHAN03(aq) + CaS04-2H20(s) (15)
The effluent gypsum slurry is pumped to a centrifuge and the resulting solids
are sent to the sludge pond. The centrate containing the dissolved ammonium
salts enters a double-effect evaporator system where the salt concentration
is increased from 15 to about 50 weight percent. This concentrated solution
can be stored and used as a fertilizer material.
Analysis of Processing Subsections
The flow diagram and material balance for the Moretana Calcium process
in the base case application are shown in Figure 4 and Table 29 respectively.
The Moretana Calcium process is a wet simultaneous SOX-NOX -system, which is
composed of 11 separate processing subsections. As an initial step, the
various pieces of process equipment were apportioned into the appropriate
subsection. Using the information available from the material balance, each
piece of equipment was described, sized, costed, and then listed by processing
subsection. This detailed equipment list for the Moretana Calcium process is
shown in Table 30. Each of these subsections is described in more detail in
the following sections.
Material Handling —
This and the following area, feed preparation, comprise the limestone
preparation section. The material-handling portion includes all of the
equipment needed to receive and unload the limestone which arrives by either
rail or truck and also the equipment needed to maintain a readily available
supply of limestone to the weigh feeders. The equipment is sized such that
only one train of processing equipment is required.
Feed Preparation —
The feed preparation area consists of a single train of equipment
required to weigh, crush, and slurry the makeup limestone. The fresh lime-
stone slurry, containing 8% solids (by weight), is temporarily stored in the
limestone slurry feed tank which has a surge volume sufficient for 8 hours
of normal operation.
Gas Handling —
Two forced-draft (FD) fans are provided between the ESP and the prescrub-
ber to overcome the pressure drop through the Moretana Calcium process. Each
fan has been sized to handle one-half of the total flue gas volume and a
pressure drop of 18.5 inches water.
79
-------
-
:
ELECTROSTATIC
6
wv
]•
/X
-J
IO
i—
=ai
II
TO
FLYA5H • 1;^^
POND
Pi / LIMESTONE
IN
' -^ \ V RAIL OR TRUCK
I
Figure 4. Moretana Calcium process. Flowsheet.
-------
TABLE 29, MORETANA CALCIUM SOX-!TOX PROCESS
MATERIAL BALANCE
stream No.
Description
|
?
\
4
=i
ti
7
H
q
10
Total stream, Ib/hr
Sft'J/Tnin (60°F>
Temoerature, °F
Pressure, psifi
Gpra
Specific gravity
pH
Undissolved solids, Z
1
Coal to boiler
428. 600
2
Combustion atr
to air heater
4,546,200
1,005,000
60
3
Combustion air
to boiler
4,101.800
906,700
535
4
Gas to
economizer
4.516,100
958 ,000
690
*,
Gas to
air heater
4.516.100
95&.000
705
Stream NO.
Description
1
/
1
4
•i
ft
7
K
4
JO
Total stream, Ib/hr
Sft3/min (60°F>
Temperature, "K
Pressure, psig
Gpm
Specific gravity
pH
llndiaaalved eoll-de, 7,
6
Gas to ESP
4,960,400
1,056,000
300
7
Tly ash
front ESP
8
Gas to
prescrubber
9
Makeup H20 to
prescrubber
332,300
10
Recycle slurry
to ptescrubber
20,429.200
• Stream No.
1
•>
J
4
"
h
7
H
cf
Total stream, Ib/hr
S£t3/nin (60°F)
Taiperature, °F
Pressure, pals
Gpm
Specific gravity
pH.
Undissolved solids, 'i
f ** 1
Preecrubber
alurry purged
to ash pond
61,000
n
Waste stream
from C102
generator
6,370
.13
Slaked lime
for pH control
7,840
14
Neutralized
effluent to
fly ash pond
125,400
127
_ 43.75
15
Effluent flue
gas from the
prescrubber
5_, 115,800
1,OT9.S15_
127
Description
t
•f
\
4
•i
fi
f
H
4
ID
Total £tream. Ib^h^
SftS^tn (60*Fi
Temperature, ^F
Gpm
Specific gravity
pH - ' '
tlndiaaolved solids, X
| is
Cl$2 gas to
flue gas ducts
53,300
ll,35i
85
IT
Oxidized flue
gag to absorber
5,169,000
1,093,129
127
_, _..lfi
Effluent
flue gas from
absorber
r,u«,ifoo
1,084, Wo
127
" ^ - -
Reheated
flue gas
to the stack
5,133,700
1.084,700
175 "~
Spent absorbent
fron absorber
64,661,200
(continued)
81
-------
TABLE 29. (continued)
1
•I
-\
4
r,
h
7
H
10
Description
Total stream, Ib/hr
Sft3/min (60°F)
Temperature, °F
Pressure, psig
Gpm
Specific gravity
pH
21
Recycle
absorbent to
absorber
64.680.000
127
1.16
5.0
27
Purge to
regeneration
section
554.300
127
1.16
5.5
2.1
Forward feed
to gypsum
centrifuge
586.700
127
1.16
5.5
9i
Wash H20
to gypsum
centrifuge
75.900
1.0
7
25
Sulfite-sulfate
sludge to
disposal pond
92.400
60
1.2
5.5
70
Stream No.
Description
1
2
(
4
S
h
7
a
2S04
to gypsum
production
reactor
8.100
60
1.76
39
Effluent
to gyspum
eentrifugp
56.800
127
1.17
7
40
Wash H20
to gypsum
centrifuge
5.000
60
1.0
7
(continued)
82
-------
TABLE 29. (continued)
Scream No.
Description
1
1
*
',
(>
7
K
1
10
Total stream. IV/hr
Sft-Vmin (60°F)
Tennerature. °F
Pressure, osls
SPefJf?c fravitv.
pH
Utidissolved solids. S
41
Gypsum to
sludge pond
11.200
-
60
20.3
1.1
7
42
Centrate to
evaporators
50.600
| 122
1.06
7
43
Fl_ld
fertilizer Co
bulk storage
16.000
1S8
1.24
^
H
_2,
10
T
•t
10
ji
<)
rT3"
83
-------
TABLE 30. MORETANA CALCIUM PROCESS
BASE CASE EQUIPMENT LIST
DESCRIPTION AND COST
Area 1 — Material Handling
1.
2.
3.
4.
5.
6.
7.
8.
9.
10.
11.
12.
Item
Car shaker and hoist
Car puller
Hopper, unloading
Feeder, unloading
vibrating
Conveyor, unloading
belt
Convey o r , unl oad ing
incline belt
Dust collector,
unloading pit
Pump, unloading pit
sump
Conveyor, storage
belt
Tripper, storage
conveyor
Mobile equipment
Hopper, reclaim
No.
1
1
1
1
1
1
1
1
1
1
1
2
Description
20 hp shaker, 7.5 hp hoist
25 hp puller, 5 hp return
16 ft dia x 10 ft straight
side height, carbon steel
3.5 hp
20 ft horizontal, 5 hp
310 ft, 50 hp
Polypropylene bag type,
2,200 ft3/min, 7.5 hp
60 gpm, 70 ft head, 5 hp
200 ft, 5 hp
30 ft3/min, 1 hp
Scraper tractor
7 ft wide x 4.25 ft high,
Total
equipment
cost, 1979 $
31,000
52,000
12,800
14,400
17,700
89,200
19,200
4,200
76,800
16,400
137,800
3,000
2 ft wide bottom, carbon
steel
13. Feeder, reclaim
vibrating
2 3.5 hp
(continued)
28,700
84
-------
TABLE 30. (continued)
14.
15.
16.
17.
18.
19.
Area
1.
2.
3.
4.
5.
Item
Conveyor, reclaim
belt
Conveyor, reclaim
incline belt
Dust collector,
reclaim pit
Pump, reclaim pit
sump
Bucket elevator,
reclaim
Bin, feed
Subtotal
2 — Feed Preparation
Item
Feeder, bin weigh
Gyratory crushers
Dust collectors,
ball mill
Ball mill
Tank, ball mill
No. Description
1 200 ft, 5 hp
1 193 ft, 40 hp
1 Polypropylene bag type
1 60 gpm, 70 ft head, 5 hp
1 90 ft high, 75 hp
2 13 ft dia x 21 ft straight
side height, covered, carbon
steel
No. Description
2 14 ft pulley centers, 2 hp
2 75 hp
2 Polypropylene bag type,
2,200 ft3/min, 7.5 hp
2 12.7 tons/hr, 172 hp
2 10 ft dia x 10 ft high,
Total
equipment
cost, 1979 $
50,700
53,800
19,200
4,200
86,400
33,200
750,700
Total
equipment
cost, 1979 $
39,600
119,500
38,400
358,500
26,800
product
5,500 gal, open top,
flakeglass-lined, carbon
steel
(continued)
85
-------
TABLE 30. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
6. Agitator, ball mill
product tank
7. Pump, ball mill
product tank
8. Tank, limestone
slurry feed
9. Agitator, limestone
slurry tank
10. Pump, limestone
slurry feed
Subtotal
10 hp, resin-lined, carbon
steel
Centrifugal belt.drive, 110
gpm, 50 ft head, 5 hp,
resin-lined, carbon steel
42.5 ft dia x 45 ft high,
477,500 gal, open top,
resin-lined, carbon steel
75 hp, resin-lined, carbon
steel
Centrifugal belt drive,
1,000 gpm, 100 ft head,
50 hp, resin-lined, carbon
steel
30,600
10,400
194,100
68,400
15,300
901.600
Area 3—Gas Handling
Item
No.
Description
Total
equipment
cost, 1979 $
1. Blower, flue gas
Subtotal
Double suction, turbofan,
3,500 hp, AP 18.5 in.
H20, carbon steel
1.472,700
1.472,700
Area 4—Particulate Control
Item
No.
Description
Total
equipment
cost, 1979 $
1. Prescrubber
Moretana tower
(continued)
893,400
86
-------
TABLE 30. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
2. Pump, prescrubber 3
recycle
3. Pump, prescrubber 2
purge
4. Tank, pH controlling 1
Centrifugal belt drive,
9,950 gpm, 75 ft head,
400 hp, resin-lined, carbon
steel
Centrifugal belt drive, 60
gpm, 75 ft head, 2 hp, resin-
lined, carbon steel
13 ft dia x 15 ft high,
14,900 gal, open top,
flakeglass-lined, carbon
steel
150,900
6,500
9,800
5.
6.
Agitator, pH tank
Pump, pH tank purge
Subtotal
1 7.5 hp, resin-lined, carbon
steel
2 Centrifugal belt drive, 150
gpm, 300 ft head, 20 hp,
resin-lined, carbon steel
13,000
17,800
1,091,400
Area 5—SCW-NOv Absorption
Item
No.
Description
Total
equipment
cost. 1979 $
1. Absorber
2. Tank absorbent
3. Pump, absorber
recycle
2 Moretana tower 1,118,300
2 45 ft dia x 50 ft high, 451,000
594,800 gal, flakeglass-lined,
concrete
3 Centrifugal belt drive, 155,400
14,000 gpm, 85.3 ft head,
700 hp, resin-lined, carbon
steel
(continued)
87
-------
TABLE 30. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
4. Pump, absorbent tank
purge
5. Agitator, absorbent
tank
6. Mist eliminator
7. Clarifier, with rake
and drive
8. Pump, clarifier
overflow
9. Tank, return liquor
10. Pump, return liquor
Subtotal
Centrifugal belt drive,
500 gpm, 50 ft head, 15 hp,
resin-lined, carbon steel
60 hp, resin-lined, carbon
steel
Euroform T-102-2, polypro-
pylene with flakeglass-lined,
carbon steel housing, 26.2
ft x 26.2 ft x 26.2 ft
45 ft dia x 10 ft high,
119,000 gal, flakeglass-
lined, carbon steel
Centrifugal belt drive, 30
gpm, 50 ft head, 1 hp, resin-
lined, carbon steel
17 ft dia x 19 ft high,
14,900 gal, open top,
flakeglass-lined, carbon
steel
Centrifugal belt drive, 250
gpm, 50 ft head, 7.5 hp,
resin-lined, carbon steel
16,700
120,100
286,100
529,700
7,000
63,400
10,800
2.758,500
Area 6—Reheat
Item
No.
Description
Total
equipment
cost, 1979 $
1. Reheater
Indirect steam, shell,
carbon steel tubes, 316L,
10,000 ft2
(continued)
1,073,000
88
-------
TABLE 30. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
2. Soot blowers
Subtotal
20 Air, fixed
130,000
1.203.000
Area 7—Gypsum Production
Item
No.
Description
Total
equipment
cost. 1979 $
1. Tank, centrifuge
feed
2. Agitator, centrifuge 2
feed tank
3. Pump, centrifuge 3
feed
4. Separator, gypsum
18
5. Conveyor, sludge to 2
mixing tank
6. Pump, gypsum
centrate to tank
7. Tank, separated
liquor
8. Pump, separated
liquor to limestone
preparation
22 ft dia x 22 ft high,
34,300 gal, open top,
flakeglass-lined, carbon
steel
15 hp, resin-lined, carbon
steel
Centrifugal belt drive, 510
gpm, 50 ft head, 15 hp,
resin-lined, carbon steel
Centrifuge, vertical basket,
resin-lined, carbon steel
Belt, 14 in. wide x 100 ft
long, 2 hp, 35 ft lift,
18.5 tons/hr
Centrifugal belt drive, 450
gpm, 50 ft head, 10 hp,
resin-lined, carbon steel
27.5 ft dia x 29 ft high,
64,800 gal, open top,
flakeglass-lined, carbon
steel
Centrifugal belt drive, 500
gpm, 100 ft head, 25 hp,
resin-lined, carbon steel
(continued)
99,400
43,100
16,700
966,600
51,800
15,900
162,200
18,200
89
-------
TABLE 30. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
9. Pump, reactor feed
10. Tank, surge
11. Pump, surge tank
recycle
Subtotal
Centrifugal belt drive, 50
gpm, 50 ft head, 1.5 hp,
resin-lined, carbon steel
45 ft dia x 48 ft high,
571,000 gal, concrete
Centrifugal belt drive,
50 gpm, 50 ft head, 1.5 hp,
resin-lined, carbon steel
7,100
330,000
7,100
1,718,100
Area 8—Nitrogen Treatment
Item
No.
Description
Total
equipment
cost, 1979 $
1. Reactor, gypsum
production
2. Agitator, gypsum
production reactor
3. Pump, reactor
effluent to
centrifuge feed
tank
4. Tank, centrifuge
feed
5. Agitator, centrifuge
feed tank
6. Pump, gypsum
centrifuge feed
15 ft dia x 17 ft high,
33,000 ft3, resin-lined,
carbon steel
10 hp, resin-lined, carbon
steel
Centrifugal belt drive, 50
gpm, 50 ft head, 1.5 hp,
resin-lined, carbon steel
10 ft dia x 11 ft high,
3,300 gal, open top,
resin-lined, carbon steel
7.5 hp, resin-lined, carbon
steel
Centrifugal belt drive, 50
gpm, 50 ft head, 1.5 hp,
resin-lined, carbon steel
(continued)
52,000
30,600
7,000
21,300
26,000
7,000
90
-------
TABLE 30. (continued)
Item
No.
Description
Total
equipment
ost, 1979 $
7. Separator, gypsum
8. Conveyor, gypsum
to mixing tank
9. Pump, centrate from
centrifuge to
evaporator feed tank
10. Pump, waste wash H20
to evaporator feed
tank
11. Tank, evaporator
feed
12. Pump to evaporators
13. Evaporator
14. Pump, fluid
fertilizer storage
feed
15. Tank, fluid
fertilizer storage
16. Pump, feed to bulk
fertilizer storage
Centrifuge, vertical
basket, resin-lined, carbon
steel
Belt, 14 in. wide x 100 ft
long, 0.5 hp, 35 ft lift,
3 tons/hr
Centrifugal belt drive, 50
gpm, 50 ft head, 1.5 hp,
resin-lined, carbon steel
Centrifugal belt drive, 100
gpm, 50 ft head, 3.0 hp,
resin-lined, carbon steel
10 ft dia x 11 ft high,
3,300 gal, open top,
resin-lined, carbon steel
Centrifugal belt drive, 50
gpm, 50 ft head, 1.5 hp,
resin-lined, carbon steel
Double effect, vacuum, 8.9
ft dia x 23 ft high,
resin-lxned, carbon steel
Centrifugal belt drive, 15
gpm, 50 ft head, 0.5 hp,
resin-lined, carbon steel
7.5 ft dia x 10 ft high,
3,300 gal, closed top,
resin-lined, carbon steel
Centrifugal belt drive, 30
gpm, 150 ft head, 3 hp,
resin-lined, carbon steel
(continued)
214,800
51,200
7,000
9,800
21,300
7,000
991,000
6,800
8,200
9,600
91
-------
TABLE 30. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
17. Tank, bulk fluid
fertilizer storage
Subtotal
35 ft dia x 40 ft high,
287,900 gal, closed top,
resin-lined, carbon steel
164,200
1,625,800
Area 9—Raw Material Storage
Item
No.
Description
Total
equipment
cost, 1979 $
1. Elevator, catalyst
to primary silo
2. Catalyst silo
3. Weigh feeder 1
4. Conveyor catalyst 1
to tank
5. Catalyst silo
6. Weigh feeder 1
7. Elevator, (NH4)2S04 1
to silo
8. Silo, (NH4)2S04
9. Conveyor, (^4)3804 2
to intermediate
surge
Bucket, 6 in. x 4 in. x 6,900
4-1/4 in., 0.25 hp, 35 ft
lift, 7 tons/hr
24 ft dia x 24 ft high, closed 71,000
top, flakeglass-lined, carbon
steel
0.5 tons/hr 6,300
Belt, 14 in. wide x 100 ft 27,100
long, 0.25 hp, 50 ft lift,
0.5 tons/hr
5 ft dia x 5 ft high, closed 1,200
top, PVC
16 Ib/hr 6,300
Bucket, 6 in. x 4 in. x 8,900
4-1/2 in., 2 hp, 50 ft lift,
13 tons/hr
40 ft dia x 42.5 ft high, 203,100
carbon steel
;Belt, 14 in. wide x 300 f.t 148,400
long, 1.5 hp, 30 ft lift,
4.1 tons/hr
(continued)
92
-------
TABLE 30. (continued)
10.
11.
12.
Item
Silo (NH4)2S04
Weigh feeder
Conveyor, (NH4>2SO^
to reactor
Subtotal
No, Description
2 10 ft dia x 13 ft high,
carbon steel
2 2 tons/hr
i 2 Belt, 14 in. wide x 100 ft
long, 1 hp, 30 ft lift,
2 tons/hr
Total
equipment
cost, 1979 $
28,300
12,600
50,000
570,100
Area 10—Sludge Disposal
Item
No.
Description
Total
equipment
cost, 1979 $
1. Tank, mixing
2. Agitator, mixing
tank
3. Pump, sludge to
pond
27.5 ft dia x 30 ft high,
133,300 gal, open top,
resin-lined, carbon steel
40 hp, resin-lined, carbon
steel
Centrifugal belt drive, 520
gpm, 300 ft head, 75 hp,
resin-lined, carbon steel
83,500
43,900
35,400
4.
5.
Sludge pond
Pump, return from
pond
Subtotal
1 253 acres
2 Centrifugal belt drive, 350
gpm, 200 ft head, 30 hp,
resin-lined, carbon steel
5,882,800
25,200
6,070,800
(continued)
93
-------
TABLE 30. (continued)
Area 11—Chlorine Dioxide Generation
Total
equipment
Item No. Description cost, 1979 $
1. Chlorine dioxide 2 22 tons/day capacity 15,709,000
generation system
Subtotal 15,709,000
Total, Areas 1-11 33,871,700
94
-------
Particulate Control—
A particulate control efficiency of 99.5% is achieved by using a 94%
efficient ESP followed by a prescrubber. (The ESP description and cost are
included later in the ESP section.) Although only a single ESP is included,
the flue gas duct is then split and two trains of fans, prescrubbers, and
absorbers are included. The prescrubber is a Moretana plate tower. The
purge stream is neutralized in pH-controlling tanks and pumped to the fly
ash pond.
SOX-NOX Absorption-
Simultaneous SOX-NOX absorption is carried out in two trains of Moretana
plate towers. Absorbent recycle tanks and pumps are also provided. In addi-
tion, the Moretana Calcium process has a separate loop between the absorber
and the reheater which includes a clarifier and mist eliminator to control
liquid carryover.
Reheat—
Flue gas reheat to 175°F is provided by two reheaters using an indirect
steam system.
Gypsum Production—
In the Moretana Calcium process no attempt is made to further treat the
mixed sulfite-sulfate salts and the mixture is simply separated, reslurried,
and pumped to the sludge pond.
Nitrogen Treatment Section—
The use of C102 as a gas-phase oxidant results in the conversion of 50%
of the absorbed NOx to nitrate salts. Since waste water nitrate salts could
result in a significant water pollution problem, a purge stream is further
treated with (NH4)2S04 to form additional gypsum sludge and ammonium nitrate
(NH4N.03) , a potential fertilizer material. The NH4N03 solution is concen-
trated in one to two double-effect evaporators and stored as a byproduct.
Raw Material Storage—
This section includes all the processing equipment needed to receive,
store, and retrieve for later use all of the various raw materials (except
limestone) needed in the process. Storage capacity is designed for 30 days
of normal usage.
Sludge Disposal—
All the equipment needed to reslurry the mixed sulfite-sulfate sludge
from the gypsum production section and the gypsum' from the nitrogen treatment
section is included. In addition, a 253-acre sludge pond is included.
C102 Generation—
Two separate C102 generating systems are provided. Each has a 22-ton/day
capacity and is designed for treating the flue gas for one absorber. Provi-
sions have been made for temporary in-process storage of C102 such that
fluctuations in the NOx concentration in the flue gas can be handled effec-
tively.
95
-------
HITACHI ZOSEN NOX-ONLY PROCESS - DRY, SCR
Process Description
Hitachi Zosen (Hitachi Shipbuilding and Engineering Company, Ltd.) has
developed an NOX FGT process for dry, SCR of NOX with NH3 (8). Hitachi Zosen
has developed a catalyst and reactor design which permits treatment of flue
gas with a high-particulate loading. Therefore, the flue gas from a coal-
fired boiler may be fed directly from the economizer to the reactor, upstream
from the air heater, without any particulate removal treatment. Before enter-
ing the reactor the flue gas is injected with NH3 which has been diluted with
air to a 5% NH3 concentration to enhance mixing. In the reactor NOx is reduced
to nitrogen by reaction with NHs in the presence of a catalyst at a tempera-
ture range of 572-752°F. Hitachi Zosen reports that the level of excess NH3
in the flue gas leaving the system is very low (<20 ppm). The treated flue
gas is passed through the boiler air heater.
The NH3 to NOX mol ratio employed for >90% removal efficiency is about
(1.0-1.2):!. Hitachi Zosen claims the pressure drop across the reactor is
2-3 inches water. The area velocity (flow rate of gas/surface area of catalyst)
is between 0.38-0.55 ft-Vmin/ft2 which is reported to correspond to a space
velocity (flow rate of gas/volume of catalyst) of 5,000-10,000 ft3/hr/ft3.
The process is reported to be capable of functioning with particulate loading
up to about 7 gr/sft3, making it possible to use this system on coal-fired
flue gas before particulate removal.
The NH3 flow rate is automatically controlled based on the flue gas rate
to the reactor, reactor inlet and outlet NOX concentration, and NH3 outlet
concentration. To prevent formation of (NH^^SO^ the catalyst bed tempera-
ture can be controlled by bypassing a part of the high-temperature flue gas
flow around the economizer to the reactor.
The. catalyst is manufactured in the shape of corrugated units. These
units are joined together for the particular size required. The flue gas
passes parallel to the. catalyst surface. The catalyst composition has not
been revealed for proprietary reasons; however, Hitachi Zosen does state
that it is constructed of common material. The catalyst life is guaranteed
for 1 year but the actual life may be higher.
Analysis of Processing Subsections
The flow diagram and material balance for the base case are shown in
Figure 5 and Table 31 respectively. The Hitachi Zosen process is divided
into three processing sections and the equipment assigned to the appropriate
section. The equipment list and descriptions by area are presented in Table
32. The total land requirement is estimated to be 1.5 acres.
NH3 Storage and Injection—
A compressor (and spare) for unloading liquid NH3 from truck or rail
transport and 250 psig storage tanks'for a 30-day supply are included.
Before NH3 is injected into the flue gas, it is vaporized in a shell-and-
96
-------
COAL
I
RAIL
TRUCK
HOOK-UP
BOILER
^ '
i
i
4
| ECONOMIZER | |
1
3
1
r 5 4 6 __
1
AMMONIA
1 STORAGE TANK
STEAM
_A *
1
F
II
111
tEACTI
1 1
i
J ,
)R _ frj AIR HEATER
I'
T
UNLOAONN
COMPRESSOR
NMj
VAPORIZER
vw
I"
Figure 5. Hitachi Zosen process. Flowsheet,
-------
TABLE 31. HITACHI ZOSSI1 HOX-011LY PROCESS
MATERIAL BALANCE
Description,
2
i
it
5
h
?
X
9
1°
Total stream, Ib/hr
Sft-»/mln <60°F)
Temperature, °F
Pressure, pslg
1
Coal to boiler
428,600
2
Combustion air
to air heater
4.546,200
1,CC5,000
80
3
Combust iori air
to boiler
4,101,800
906,700
535
4
Gas to
economizer
4,516,100
95&,UUU
890
5
Flue gas
to reactor
4,516.100
95&,QOO
705
S tieam T3o.
1
2
)
l<
•,
h
1
a
9
IV
Description
Total stream, Ib/hr
SftJ/min (60°F)
Temperature, °F
Pressure, pslg
6
Flue gas-NHj
mixture to
reactor
4,576,700
971,800
697
-r
Gas to air
heater
4,576,700
971,800
712
8
Gas to ESP-
5,021,000
1,070,200
300
9
Gas to FGD unit
4,966,500
1,070,200
300
10
Treated £lue
eas to st^ck
5,165,600
1,143,400
175
Stream No.
1
*
4
^
h
7
K
&
IK
TJescr lotion
Total stream, Ib/hr
Sft3/min (60°F)
Temperature, ^F
Pressure, psig
11
HH3 scream
injected into
flue gas
60.60p
.. 13. 600
12
NHj from
storage
1.827
110
225
13
Steam to
vaporizer
1,120
298
50
14
Fly ash
£rom ESP
54.600
a
9
10
98
-------
TABLE 32. HITACHI ZOSEN NOX-ONLY PROCESS
BASE CASE EQUIPMENT LIST
DESCRIPTION AND COST
Area 1—NH3 Storage and Injection
Item
No.
Description
Total
equipment
cost, 1979 $
1. Compressor, NH3
unloading
2. Tank, NH3 storage
3. Vaporizer, NH3
4. Blower, NH3 and
air
Subtotal
Single cylinder, double
acting, 300 sft3/min at 250
psig, 30 psig suction, 125
hp, cast iron
Horizontal, 9 ft dia x 66
ft long, 30,000 gal, 250
psig, carbon steel
Steam at 298°F, tube type,
30 ft2, 1.03 MBtu/hr, carbon
steel
7,000 aft3/min, AP 5 in. H20,
10 hp, carbon steel
50,000
230,100
36,800
13,500
330,400
Area 2—Reactor Section
Item
No.
Description
Total
equipment
cost, 1979 $
1. Reactor
2. Initial catalyst
fill
Subtotal
40 ft x 40 ft x 31 ft high,
715°F operating temperature,
carbon steel, insulated, with
fly ash hoppers and sootblowers
Proprietary catalyst, manu-
factured in honeycomb-shaped
units
1,029,000
6,370,000
7.399.000
(continued)
99
-------
Area 3—Flue Gas Fans
TABLE 32. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
1. Blower, flue gas
Subtotal
Total, Areas 1-3
391,025 aft3/min, AP 22 in.
H20, 2,250 hp, 316 stainless
steel
405,000
405,000
8,134,400
100
-------
tube steam-heated vaporizer and mixed with air supplied by a small blower
to form a 5% NH3-in-air mixture, since the inflammability limits for NH3 in
air are 15.5 to 27.0%, and for easier flow control and for mixing enhancement.
There is one vaporizer and one air-NH3 blower for each of the two reactor
trains. The assumed pressure drop for sizing the NH3~air blower is 5 inches
water.
The NH3 injection grid is primarily piping and the cost therefore is
incorporated into the area investment estimate shown in the summary capital
investment estimates in the following report section, Economic Evaluation and
Comparison.
Reactor Section—
Two reactors are provided, each handling 50% of the total flow. The
reactors are fixed-bed type and constructed of carbon steel. Each reactor
is provided with fly ash hoppers for collection of any deposited fly ash arid
with two soot blowers for periodic cleaning. The catalyst is manufactured
into corrugated shaped units and these units are Joined together to form the
appropriately sized catalyst bed.
Gas Handling—
A larger ID blower is provided downstream of each ESP to compensate for
the pressure drop created by the boiler, FGT system, and ESP. However, only
the incremental cost attributed to the NOx removal system is included in the
cost estimates. Because of the larger pressure drop, additional costs to prevent
implosion are included with gas handling.
KURABO KNORCA NOX-ONLY PROCESS - DRY, SCR
Process Description
The Kurabo Knorca process is based on SCR of NOX with NH3 using a moving-
bed catalyst developed by Kurabo (5). Two types of catalyst systems have
been developed by Kurabo. One system uses copper oxide on an aluminum oxide
carrier which requires thermal regeneration after each pass through the reac-
tor. The other system is an iron-oxide catalyst on a titanium dioxide carrier
which, according to the process developers, does not require thermal regenera-
tion and is more applicable to treatment of coal-fired flue gas and for gases
in which chlorides exceed 100 ppm. The latter catalyst is used for this
evaluation.
The Knorca reactor removes about 80% of the fly ash in the flue gas but it
cannot be operated at the full-fly ash loading of coal-fired flue gas.
Therefore, a 98% efficient hot ESP is used between the economizer and reactor
to reduce the fly ash level in the flue gas going to the reactor to 0.44
gr/sft-*. The overall particulate removal efficiency of the system is 99.5%.
The flue gas leaving the ESP is injected with NH3 and passed through a static
gas mixer to disperse the NH3 in the flue gas before it enters the reactor.
The reactor contains a bed in which the spherical catalyst <3-5 mm diameter)
moves slowly and continuously downward as the flue gas passes through it in
crossflow. The NOx is reduced to molecular nitrogen and water in the presence
101
-------
of the catalyst. The flue gas leaving the reactor continues through the
boiler air heater. Fly ash and broken catalyst are separated from the cata-
lyst removed from the reactor by a vibrating sieve, and the catalyst is
conveyed by horizontal vibrating conveyors and bucket elevators to the
reactor feed hopper. Makeup catalyst is manually added to replace the
broken catalyst. A baghouse and air blower are used to collect the fly
ash separated from the catalyst in the vibrating sieve.
The reactor design operating conditions are 716°F, atmospheric pressure,
and a space velocity of 7,500 ft3/hr/ft^. The flue gas residence time is
about 0.15 second in the catalyst bed and 2-3 seconds in the reactor. The
pressure drop across the reactor and static gas mixer is 4 inches water and
2 inches water respectively. The catalyst panels in each reactor are divided
into modules of 10- to 15-MW individual capacity. The expected catalyst life
is at least 2 years and the loss makeup rate is about 10%/year.
The NH3 flow rate is automatically controlled on the basis of the inlet
gas NOX concentration and flue gas rate. As an additional control, the NH3
usage rate is also compared with fuel consumption. Flue gas temperature can
be controlled by passing a part of the flue gas stream around the economizer.
TRW, Inc., has exclusive rights to supply the Knorca system for boilers
in the United States and Canada.
Analysis of Processing Subsections
The flow diagram and material balance are shown in Figure 6 and Table 33
respectively. The Knorca process is divided into four processing sections and
the equipment assigned to the appropriate section. Table 34 shows the equip-
ment list and description by area. The total land requirement is estimated
to be 1.5 acres.
NH^ Storage and Injection—
A compressor (including a spare) for unloading liquid NH3 from truck or
rail transport and 250 psig storage tanks sufficient to maintain a 30-day
supply are included. Four steam-heated, shell-and-tube vaporizers are pro-
vided, one for each reactor train. The NH3 injection grid is primarily piping
and the cost is therefore incorporated into the area investment estimates
shown in the following report section, Economic Evaluation and Comparison.
Reactor Section—
A static gas mixer is used)to ensure thorough dispersion of the NH3 in
the flue gas prior to the mixture entering the reactor. There are four
reactor trains, as recommended by Kurabo, and each is equipped with a cata-
lyst processing system. Each reactor contains 92.6 tons of the iron-based
catalyst. The catalyst panels are divided into modules of 10- to 15-MW
equivalent capacity. The reactor housing is carbon steel, but the screen
which interfaces the flue gas and catalyst is stainless steel. The sieve
which separates fly ash and broken catalyst from the catalyst is carbon steel
construction. The recycled catalyst is transported to a bucket elevator by
a horizontal vibratory conveyor sized for 0.53 ton/hr. The catalyst is
102
-------
H
s
RAIL
°* r »JH .
TRUCK e
HOOK-UP **
UNLOADING
COMPRESSOR
BUCKET
VIBRATORY tUEVATOT
FEEMI
BROKEN
CATALTST
CATALYST
HAKE - UP
| VIBRATORY HORIZONTAL CONVITOR
8*6
HOUSE
Figure 6. Kurabo Knorca process. Flowsheet.
-------
TABLE 33. KURABO KNORCA HOX-OULY PROCESS
MATERIAL BALANCE
2
t
f)
8
9
!».
Total stream, Ib/hr
Sft3/ffiin (60°F)
Temperature, °F
Pressure, psig
1
428,600
2
Combustion air
4,546,200
1,005,000
80
3
Combustion air
to boiler
4,101,800
906,700
535
4
Gas to
economizer
4.516.100
958,000
890
Gas to ESP
4,516,100
958,000
705
1
•>
!
8
q
10.
Total stream, Ib/hr
Sft^/mln (600F)
Pressure, palg
" ~~
6
Flue gas to
4,462,300
958,000
705
7
Flue gas-NH3
mixture to
4,463,800
958,600
705
I 5 1
Gas to air
heater
4.462.900
958.700
717
Gas to FGD unit
4.907.300
1,057.000
300
10
Treated flue
gas to stack
1,130,000
175
Description
1
•>
1
4
h
7
H
9
Total stream, Ib/hr
Sft3/mln (60°F)
Pressure, psig
11
NH3 from
storage
1.533
570
110
225
12
Steam for NH3
vaporization
940
298
50
13
Catalyst and
fly ash to
sieve
5.110
14
Broken
catalyst
10.58
"
Circulating
catalyst
4,222
Stream No.
Description
1
;
1
h
;
H
y
10
Total stream, Ib/hr
SftVmin C60°F)
Temperature, °F
16
Catalyst
10.58
17
Reactor
4.233
18
Gas stream
to baehouse
29,051
6,280
19
Fly ash exiting
baghouse
878
20
Air from
baahouse
28.200
5.850
100
Atm
(continued)
104
-------
TABLE 33. (continued)
Strean^ N9T
Description
I
j
i
A
')
&
7
K
t
1°
Total stream. Ib/hr
Sft3/mln (60"F)
Temcerature. °F
Pressure, palE
21
ny ash
from ESP
53.800
I
2
~Ti
i4
~
~T
To"
105
-------
TABLE 34. KURABO KNORCA NOX-ONLY PROCESS
BASE CASE EQUIPMENT LIST
DESCRIPTION AND COST
Area 1—NH3 Storage and Injection
Item
No,
Description
Total
equipment
cost, 1979 $
1. Compressor, NH3
unloading
2. Tank, NH3
storage
3. Vaporizer, NH3
Subtotal
Single cylinder, double
acting, 300 sft^/min at
250 psig, 30 psig suction,
125 hp, cast iron
Horizontal, 9 ft dia x 66
ft long, 30,000 gal, 250
psig, carbon steel
Steam at 298°F, tube type,
15 ft2, 1.03 MBtu/hr,
carbon steel
50,000
230,100
52,200
332.300
Area 2—Reactor Section
Item
No.
Description
Total
equipment
cost, 1979 $
1. Static gas mixer
2. Reactor
3. Feeder, reactor
to sieve
13.5 ft x 13.5 ft, minimum
length 25 ft, 35,800 aft3/
sec at 705°F, carbon steel
21.3 ft wide x 82 ft long x
42.7 ft high, 715°F
operating temperature,
atmospheric pressure, space
velocity = 7,500 hr"1,
housing is carbon steel,
screen is stainless steel
Vibratory, 0.64 tons/hr,
21.8 ft3/hr, 0.25 hp
(continued)
86,400
1,996,800
5,600
106
-------
TABLE 34. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
4. Sieve
5. Conveyor, sieve
to bucket
elevator
6. Elevator, circu-
lating catalyst
to catalyst
reservoir
7. Conveyor, bucket
elevator to
catalyst hopper
8. Hopper, catalyst
0.64 tons/hr, 21.8 ft3/hr,
1 hp, screen and connector
are stainless steel
Vibratory horizontal, 82
ft long (max), 0.53 tons/
hr, 1 hp, trough =9 in. x
2 in.
Continuous, bucket = 6 in. x
4 in., 130 ft high (max),
0.53 tons/hr, 17 ft3/hr, 2 hp
Vibratory horizontal, 82 ft
long (max), 0.53 tons/hr,
1 hp, trough = 8 in. x 2 in.
7.5 ft dia x 8 ft high, 353
ft3 (without conical bottom)
conical bottom = 6.5 ft high,
449 ft3 total, carbon steel
20,000
19,400
48,400
19,400
23,500
9. Initial catalyst
"fill
Subtotal
4 92.6 tons/train, Fe-based 3,936,300
catalyst on Ti02 support
6,155,800
Area 3—Waste Disposal
Item
No.
Description
Total
equipment
cost, 1979 $
1. Baghouse filter
1,666 aft3/min, 220 Ib/hr,
installed cloth area = 833
ft^, operating cloth area *
666 ft2
(continued)
32,000
107
-------
TABLE 34. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
2. Blower, baghouse
Subtotal
1,666 aft3/min, AP 4 in.
H20, 2 hp, carbon steel
13,400
45,400
Area 4—Flue Gas Fans
Item
No.
Description
Total
equipment
cost, 1979 $
1. Blower, flue gas
Subtotal
386,000 aft3/min, AP 22
in. H20, 2,250 hp, 316
stainless steel
405,000
405,000
Total, Areas 1-4
6,938,500
108
-------
transporated by another vibratory conveyor from the bucket elevator to the
catalyst hopper which feeds the reactor. The catalyst hopper is carbon steel
and has a 449 ft-* volume.
Waste Disposal—
The fly ash removed from the catalyst at the sieve is collected by a
baghouse filter for disposal. The baghouse filter and baghouse blower are
sized on the basis of: (1) fly ash loading of 40 gm/Nm3, (2) operating air/
cloth ratio of 2.5 aft3/min/ft2 cloth, (3) installed air/cloth ratio of 2.0
aft3/min/ft2 cloth, and (4) pressure drop of 4 inches water. There is one
baghouse filter and blower per reactor train.
Gas Handling—
Four higher energy ID blowers are provided downstream from the air
heater to compensate for the pressure drop created by the NOx removal system.
Only the incremental costs attributed to the increased pressure drop of the
NOX removal system are included in the estimates. A pressure drop of 7 inches
water for the NOX removal system is used as a basis. Because of the larger
pressure drop, additional costs to prevent implosion are included with gas
handling.
UOP SFGT-N, NOX-ONLY PROCESS - DRY, SCR
Process Description
In addition to simultaneous SOX-NOX removal the UOP process can be used
as an N0x-only process (21). The same catalyst—copper sulfate on an aluminum
oxide carrier—as used in the S0x-N0x process is used but the catalyst is not
regenerated. When the copper oxide is converted to copper sulfate by SOX in
the flue gas, it remains in this form and the SOX passes unchanged through the
system. Flue gas leaving the economizer is injected with an NH3-air mixture
and passed through a fixed-bed, parallel-passage reactor where NOx is reduced
to molecular nitrogen and water by the NH3 in the presence of the catalyst.
The NH3 is diluted with air. to a 5% by volume NH3-air mixture before injection
to ensure adequate mixing with the flue gas.
The reactor operating conditions are similar to those of the UOP SFGT-SN
SOX-NOX removal process:
Maximum particulate loading Full loading (>10 gr/sft3)
Pressure drop across the reactor 5-6 inches water
NH3:NOX mol ratio (1.0-1.2):!
Average space velocity 4,000-8,000 ft3/hr/ft2
Temperature 750°F
With the NOX FGT system located between the economizer and air heater,
the recommended method of controlling flue gas reaction temperature consists
of regulating a bypass flow of either flue gas or boiler feedwater around the
economizer to control the heat transfer rate. Also, all the reactions occur-
ring in the reactor are exothermic and by having the FGT system located
upstream of the air heater this heat of reaction can be recovered from the
flue gas.
109
-------
The parallel-passage reactor has been shown to be capable of operating
at normal conditions with full particulate loading (>10 gr/sft^), although
reactor cleaning methods were necessary in some cases. Initial testing with
full-fly ash loading to a pilot-plant reactor at a test unit in Pernis, The
Netherlands, showed no deterioration in performance. However, at a 40-MW
equivalent facility at Showa Yokkaichi Sekiyo refinery in Yokkaichi, Japan,
the high vanadium and sodium content of the fly ash gradually fouled the
reactor internals, limiting operating periods to 1-2 months. Modifications
recently have allowed an operating period of approximately 1 year without
significant deterioration in performance. The different particulate loading
with coal-fired flue gas (>10 gr/sft3) at a test facility on a Tampa Electric
Company boiler caused a decline in performance after only days of operation.
As a result of this problem, a procedure was implemented at this pilot plant
which provided in situ cleaning of reactor internals during normal operation.
This technique allowed stable performance at high-particulate loadings without
any adverse effects on the reactor pressure drop or on the FGD system.
Shell International Petroleum Maatschappij developed this process and
has been involved in the program since the early 1960's. The UOP Process
Division holds the worldwide licensing rights to the SFGT process (with the
exception of the Far East).
Analysis of Processing Subsections
The flow diagram and material balance are shown in Figure 7 and Table 35
respectively. The UOP SFGT-N process is divided into three sections and the
process equipment assigned to the appropriate section. Table 36 shows the
equipment list and description by area. The total land requirement is
estimated to be 1.5 acres.
HH3 Storage and Injection-—
A compressor, including a spare, for unloading liquid NH3 from truck or
rail transport and storage tanks for a 30-day supply of NH3 at 250 psig are
included. A steam-heated, shell-and-tube vaporizer is provided for vaporiza-
tion of NH3, prior to its being diluted with air. The air for dilution of the
NH3 is supplied by a blower sized for a 5-inch water pressure drop. There is
a vaporizer and blower for each of the two reactor trains. The NHo injection
grid is primarily piping and the cost is incorporated into the area investment
estimate shown in the summary capital investment estimates in the following
report section, Economic Evaluation and Comparison.
Reactor Section—
Two reactors are provided, each of which handles 50% of the flue gas and
NH3 mixture. The reactors are of parallel-passage design and made of carbon
steel. The catalyst is contained within unit cells and the flue gas is forced
across the face of the catalyst layers, not through the catalyst.
110
-------
COAL
I
BOILER
ECONOMIZER
TRUCK ^
HOOK-UP
AMMONIA
STORAGE TANK
STEAM
1(3
UNLOADING
COMPRESSOR
12
NMj
VAPORIZER
FACTOR
AIR HEATER
AIR
ELECTROSTATIC
PRECIPITATOR
Figure 7. UOP SFGT-N, N0x-only process. Flowsheet
-------
TABLE 35. UOP SFGT-N, NQX-ONLY PROCESS
MATERIAL BALANCE
Stream No.
Description
!
j
1
;
•,
f>
,-
K
')
Id
Total scream. Ib/iir
Sft3/mln (60°F)
Temperature. °F
Pressure^ pslK
1
Coal to boiler
428.600
2
Combustion air
to air heater
'..546.200
1 ,005.000
80
3
Combustion air
to boiler
4.101.800
906,700
515
i,
Gas to
economizer
4. sift, inn
958,000
890
5 >
!
Flue gas to
reactor
4,516,100
958,000
705
Stream Ko.
Description
t
1
1
•
>
'»
;
rt
'I
iu
Total stream. Ib/hr
Sfti/nin (60°F)
Temperature, °F
Pressure, psiR
f>
Flue gas-NH3
mixture to
reactor
4,576 70fl
971.800
597
1
Gas to air
heater
4,5?6,70Q
971,800
712
8
Kas to ESP
5,021,000
1,070,200
300
9
Gas to FCD unit
4,966,400
1,070, 200
300
10
Treated flue
pas to stack
5,168,700
1,143,700 j
175
.
j
Stream No,
Description
1
i
i
i
h
.
N
4
II!
Total stream. Ib/hr
Sft3/min f60°F)
Temperature, °F
Pressure, psig
11
NH^ stream
injected Into
flue Ras
60.600
13.800
12
NH3 from
storage
1,827
110
225
13
Steam to
vaporizer
1,120
298
50
14
Fly ash
from ESP
54,600
112
-------
TABLE 36. UOP SFGT-N, NOX-ONLY PROCESS
BASE CASE EQUIPMENT LIST
DESCRIPTION AND COST
Area 1—NH3 Storage and Injection
Item
No.
Description
Total
equipment
cost, 1979 $
1. Compressor, NH3
unloading
Tank, NH3
storage
3. Vaporizer, NH3
2 Single cylinder, double
acting, 300 sft3/min at
250 psig, 30 psig suction,
125 hp, cast iron
8 Horizontal, 9 ft dia x 66 ft
long, 30,000 gal, 250 psig,
carbon steel
2 Steam at 298°F, tube type, 30
ft2, 1.03 MBtu/hr, carbon
steel
50,000
230,100
36,800
4.
Blower, NH3 and air
Subtotal
2 7,000 aftVmin, AP, 5 in.
H2.0, 10 hp, carbon steel
13,500
330,400
Area 2—Reactor Section
Item
No.
_D_e_s cr Ip t ion
Total
equipment
cost. 1979 $
1. Reactor
39 ft dia x 24 ft long, 715°F
operating temperature, pressure
same as flue gas duct, carbon
steel, insulated
5,428.000
Included is a prorated allowance
for a UOP-SFGT package consisting
of reactor internals with initial
catalyst fill, process engineering,
license fees, and startup assistance
Subtotal
5.428.000
(continued)
113
-------
TABLE 36. (continued)
Area 3—Flue Gas Fans
Total
equipment
Item No. Description cost. 1979 $
1. Blower, flue gas 4 391,027 aft3/min, AP 22 in. 405,000
H20, 2,250 hp, 316 stainless
steel
Subtotal 405.000
Total, Areas 1-3 6,163,400
114
-------
Gas Handling —
A larger ID blower is provided downstream of each ESP to compensate for
the pressure drop created by the boiler, FGT system, and ESP. However, only
the incremental cost attributed to the NOX removal system is included in the
cost estimates. A pressure drop of 7 inches water for the NOX process is
used as a basis for the incremental cost calculations. Because of the larger
pressure drop, additional costs to prevent implosion are included with gas
handling.
UOP SFGT-SN, SOX-NOX PROCESS - DRY, SCR
Process Description
UOP offers a dry, simultaneous SOX-NOX removal system using the Shell CuO
H2S04.
The flue gas leaving the boiler economizer is injected with NH3 and is
sr -
T
the flue gas by the following reaction:
The S02 in the flue gas reacts with this sorbent as shown below:
S02(g) + l/202(g) + CuO(g) - CuS04(s) (2)
^ XTA 4n HIP flue Kas is reduced to molecular
In the presence of this CuSOA the NOx in the ri g reactor and
nitrogen and water by NH3. The "eated flue gas th«n^ ^ ^ ^
passes through the steam superheater, trie DOIA
efficient "cold" ESP before going to the stacK.
When the acceptor material becomes J?*"™ f iis diverted t
efficiency drops below 90% removal, ^e tiu^g ^ acceptor is regenerated
containing regenerated acceptor »«eri • neratlon gas which is produced
with a steam-diluted, hydrogen-containing^ gources of hydrogen which could
by steam-reforming naphtha. ^Jere. \ since this regeneration occurs at
be considered, e.g., coal gasiflca C1°"'; or cooiing of the acceptor is not
the same temperature as acceptance, n *tal r (Cu) , S02, and water,
required. The CuSOA is convert ed to «£* ter> The following equations
while any unreacted CuO Is reduced to Cu and wate
express the above reactions:
CuSOA(s) + 2H20(g) - Cu(g) + S02(g) + 2H20(g) (3)
Cu°(s) + H2(g) * C
115
-------
The off-gas produced during regeneration contains S02, water vapor, and traces
of unreacted reducing gas and may be handled by various means depending on
the desired product and the inerts in^ the reducing gas. For this study this
off-gas is sent to a captive sulfuric acid plant to produce acid. (An alter-
native S02 processing scheme which could be considered is the generation of
elemental sulfur from a Glaus unit,)
Since this FGT process is cyclic, the off-gas production is also cyclic
and before further processing of the regeneration off-gas flow variations
must be minimized. The bulk of the water vapor also must be condensed.
Therefore, the off-gas passes through a waste heat boiler to recover sensible
heat and a direct-contact cooler for cooling below the dewpoint. A gas
compressor and gas holder are considered the simplest system for dampening
the flow.
Normal operating conditions for the UOP SFGT-SN process are as follows:
Maximum particulate loading Full loading (-10 gr/sft^)
Pressure drop across the reactor 15-16 inches water
NH3:NOX mol ratio (1.0-1.2):!
Average space velocity 5,000-8,000 ft3/hr/ft2
Temperature 750°F
Under these conditions both NOX and SOX removal efficiencies will be 90%.
Due to the unique processing scheme of the UOP SFGT-SN process, various
additional benefits can be expected, particularly in the total energy consump-
tion. Since the reactions occurring in the acceptance cycle are exothermic
and the FGT system is located upstream of the air heater, the various heats
of reaction are recoverable in the boiler air heater. The sensible heat of
the regeneration off-gas is also recovered in the waste heat boiler and
contact cooler. The process vendors also indicate that since 803 is removed
by the sorbent, the dewpoint is lowered and therefore more sensible heat is
potentially recoverable in the boiler air heater.
Shell International Petroleum Company developed this process and has ;
been involved in the program since the early 1960's. The UOP Process Division
holds the worldwide licensing rights to the process (with the exception of
the Far East).
Analysis of Processing Subsections
The flow diagram and material balance for the base case are shown in
Figure 8 and Table 37 respectively. The process is divided into six processing
sections and the equipment assigned to the appropriate sections. The equip-
ment list and descripions are shown in Table 38 by area.. The total land
requirement for this base case is estimated to be 4 acres.
116
-------
BOILER
4 _
ECONO
COAL
1
t7
UNLOADING
RAIL COMPRESSOR
OR i ,| 0 ,
TRUCK ' TST^
HOOK-UP
1 AMMONIA \
1 STORAGE TANK 1
"i""
*H, L-^-S-
VAPORIZER J
MIZER 0 ^
S
BYPASS
t
f r " '
^
f" u "T ^_L^_^ si»
•iliTT
r— r
REACTOR REACTOR
1 r^P
STEAM f
'ERHEATER |Z
AIR
TO
, — • STEAM
HEADER
/STEAM DISENGAGING^
^ DRUM )
i
STEAM
NAPHTHA
REFORMER
9 ELECTI
" PRECIf
\A
/
STACK
10 G '
IOSTATIC
ITATOR
A/
i-
N^ FLUE GAS
>^ TO
ATMOSPHtRE
. t •
o 6 • - J
BOILER 53 36
• ^^/-' •- .••,!;
WASTE HEAT
BOILER
17
t
20
-a
SUCTION DRUM
2 QUENCH
TOWER
2»
n«-?fATmMER QUENCH WATER
^-,— ^ p"h?J COOLER
El U U
5TRIPPES BOTTOMS
COOLER
i-r — ' *• — ' *• — Tii
35 ,
CC
] FILTER
COMPRESSOR
28
STRIPPER
31
90 i,
FILTER SSTRIPPER BOTTOMS
CONDENSATE D U PUMP
X— ~>v
it PLANT
MPRESSOR ^— L^
QUENCH
TOWER 5^ K
HOLDING
. DRUM
SULFURIC
^-^ \^_J ACID
21 20
NOTE DASHED LINES REPRESENT ALTERNATE CYCLES
Figure 8. UOP SFGT-SB, SOX-NOX process. Flowsheet.
-------
TABLE 37. UOP SFG'i'-SN, SOX-1IOX PROCESS
MATERIAL BALANCE
Description
1
1
1
V
>
i.
1
«
9
IO
Total stream, Ib/hr
SftJ/min <60°F)
Temperature, °F
Pressure, psig
1
Coal to boiler
428,600
2
Combustion air
to air heater
4, 546. 200
1^005,000
80
3
Combustion air
to boiler
A, 101, 800
906,700
535
4
Gas to
economizer
4.516.100
958,000
S90
->
Gas to blower
4,516.100
958,000
705
Stream No.
Description
1
>
!
'.
rt
(S
!
,S
y
l(?
Total stream, Ib/hr
Sft3/mln (60"P)
Temperature, °F
Pressure, psig
f,
Flue gas-NH3
mixture to
reactor
4,576,700
971,800
720
7
Flue gas
to s t earn
superheater
4,548,800
969,900
798
8
Gas to air
heater
4,548,800
969,900
786
9
Gas to ESP
4,993,200
1,068,100
300
10
Gas to stack
4,938,600
1,068,100
300
Stream No.
Description
1
J
f
--,
•>
f>
/
H
9
10
Total stream, Ib/hr
SftJ7iin (60°F)
Temperature, °F
Pressure, psig
11
NHj stream
injected Into
flue gas
60,600
13,800
12
Steam to steam
superheater
65.500
298
50
13
Hydrogen for
regeneration
2.870
5,700
14
Steam to purge
6,270
650
12
15
Reducing gas
to reactor
62,100
26,200
Stream No.
1
2
I
'<
>
6
/
H
V
10
Description
Total stream, Ib/hr
SftJ;rain (60°F)
Temperature, °F
Pressure, psiR
16
Regeneration
gas to waste
heat boiler
96.250
28,100
17
Regeneration
gas to quench
tower
96.250
28,100
18
Boiler
feedwater to
steam
disengaging
drum
12.710
113
180
19
Steam to
steam header
12.710
36T
150
20
Compressor
quench tuwrr
bottoms to j
quench tower :
28r520
190 .
(canttnued)
118
-------
TABLE 37. (continued)
Stream No,
Description
i
j
-*
^
h
7
H
9
10
Total stream, Ib/hr
Sft'3/mln (60°F)
Temperature. °F
Pressure, pslg
21
Stripper
overhead to
quench tower ^
360
190
4.5
22
Quench tower
feed
125,330
23
Lower quench
tower recycle
663,970
113
125
24
Upper quench
touer recycle
663.970
113
125
25
Quench tower
overhead
25,400
m
5
Stream No.
1
2
!
4
rt
l>
!
8
9
11?
Description
Total stream, Ib/hr
Sft3/min (60°F)
Temperature, °F
Pressure, pslg
26
Main quench
tower bottoms
1,427.870
190
27
Main quench
bottoms to
cooler
1.355,570
190
150
28
Main quench
bottoms from
cooler
1,355.570
113
125
29
Main quench
bottoms to
stripper
72,300
190
135
30 /•
i,
Stripper bottoms
to cooler
76,590 J
C
241
70
Stream No.
1
J
4
')
6
7
H
Temperature, °F
Pressure, psig
31
Steam to
stripper
4.850
298
50
32
Stripper
bottoms from
cooler
76.590
113
65
33
Compressor
suction
34.880
122
4
34
Compressor
discharge
34.880
356
65
35
Compressor
quench tower
spray
27,630
J
113
125 t
Stream No.
1
J
1
4
5
h
1
H
-------
TABLE 37. (continued)
Description
1
1
it
r
h
H
9
12-
Total stream, Ib/hr
Sft-Vmin (6QOF)
Temperature, °F
Pressure, psig
41
Purge gas to
reactors in
acceptance
stage
8.160
2,900
42
NH3 from
storage
1,827
110
225
Steam to
vaporizer
1,120
298
50
Fly ash
from ESP
8
9
LLOJ
9
10
_H
9
10
120
-------
TABLE 38. UOP SFGT-SN, SOX-NOX PROCESS
BASE CASE EQUIPMENT LIST
DESCRIPTION AND COST
Area 1—NH3 Storage and Injection
Item
No.
Description
Total
equipment
cost. 1979 $
1. Compressor, NH3
unloading
2. Tank, NH3 storage
3. Vaporizer, NH3
N Blower, NH3 and air
Subtotal
Single cylinder, double
acting, 300 sft3/min at
250 psig, 30 psig suction,
125 hp, cast iron
Horizontal, 9 ft dia x 66 ft
long, 30,000 gal, 250 psig,
carbon steel
Steam at 298°F, tube type,
30 ft2, 1.03 MBtu/hr,
carbon steel
7,000 aft3/min, AP 5 in.
H20, 10 hp, carbon steel
50,000
230,100
36,800
13,500
330,400
Area 2—Flue Gas Fans
Item
No.
Description
Total
equipment
cost. 1979 $
1. Blower, flue gas
Subtotal
1,088,400 aft3/min, AP 17.1
in. H20, 4,000 hp, 316
stainless steel
1.510,300
1,510,300
(continued)
121
-------
TABLE 38. (continued)
Area 3—Reactor Section
Item
No.
Description
Total
equipment
cost, 1979 $
1. Reactors
2. Steam super-
heater
Subtotal
23.5 ft dia x 30 ft long, 16,969,000
design temperatures -
1,060°F top, 800°F bottom,
22 psig internal, 1.37
psig external, carbon
steel, insulated
Included is a prorated allow-
ance for a UOP-SFGT package
consisting of reactor internals
with initial catalyst fill,
flue gas valves with hydraulic
turning gear (2/reactor),
process engineering, license
fees, startup assistance,
and for incremental air
heater investment
Steam at 298°F, tube type, 267,500
3,000 ft2, 17.1 MBtu/hr,
carbon steel
17,236,500
Area 4—Flow Smoothing
Item
No.
Description
Total
equipment
cost, 1979 $
1. Boiler, waste heat
2. Steam disengaging
drum
Boiler feedwater at 113°F,
150 ft2, 19.2 MBtu/hr,
tube = 25 psig, shell =
200 psig, carbon steel
9 ft dia x 25 ft long,
420°F, 200 psig, carbon
steel, insulated
29,800
38,100
(continued)
122
-------
TABLE 38. (continued)
Item
No.
Description
Total
equipment
cost, 1979 $
3. Steam blowdown
drum
4. Main quench tower
5. Pump, quench water
circulating
6. Filter, quench
feed
7. Cooler, quench
water
8. "Stripper
Pump, stripper
bottoms
10. Filter, stripper
bottoms
11. Cooler, stripper
bottoms
Vertical, 5 ft dia x 15 ft
long, 420°F, atmospheric
pressure, carbon steel,
insulated
6 ft 6 in. top dia, 13 ft
bottom dia, 40 ft long,
420°F, 50 psig, 317L
stainless steel clad
Centrifugal, 3,100 gpm, 368
ft head, 300 hp, casing,
317L stainless steel;
impeller, 317L stainless
steel
65 ft2 of cotton media,
Dollinger model
Air cooled fin fan type,
104.3 MBtu/hr, 23,600
110 psig, 317L stainless
steel
6 ft dia x 20 ft long,
280°F, 50 psig, 317L
stainlers steel clad
Centrifugal, 155 gpm, 172 ft
head, 15 hp, casing, 317L
stainless steel; impeller,
317L stainless steel
38 ft2 of cotton media,
Dollinger model
Air cooled fin fan type,
9.96 MBtu/hr, 1,800 ft2,
90 psig, carbon steel
(continued)
13,700
189,600
110,700
30,000
793,100
51,100
16,000
15,000
50,000
123
-------
TABLE 38. (continued)
Item
No.
Description
Total
equipment
cost. 1979 $
12. Compressor suction 1
drum
13. Compressor, regenera- 1
tion rich gas
14. Compressor, quench
tower
15. Gas holding drum
Subtotal
5 ft dia x 10 ft long,
175°F, 50 psig, 317L
stainless steel clad
4,300 aft3/min at 122°F,
suction temperature, 122°F;
suction pressure, 967 mm Hg
abs; discharge pressure,
4,121 mm Hg abs, 550 hp,
900 rpiii (max) , 317L stain-
less steel
4 ft dia x 14 ft long,
410°F, 90 psig, 317L
stainless steel clad,
insulated
12 ft dia x 24 ft long,
410°F, 90 psig, 317L
stainless steel clad,
insulated
41,800
100,000
68,100
127.000
1,674.000
Area 5—Steam Naphtha Reformer
Item
No.
Description
Total
equipment
cost, 1979 $
1. Packaged steam
naphtha reformer
Subtotal
1 10 Msft3/day, 1,500 gal/hr
naphtha, plot area = 50 ft x
75 ft, skid mounted
(continued)
3.222.000
3.222.000
124
-------
TABLE 38. (continued)
Area 6—H2S04 Plant
Item
No.
Description
Total
equipment
cost, 1979 $
1. H2S04 plant,
contact unit
Subtotal
Total, Areas 1-6
427 tons/day, 100% H2S04
as 98% or 93% acid, flue
gas meets EPA standards
5.719,000
5,719.000
29,692,200
125
-------
NH3 Storage and Injection—
A compressor (including a spare) for unloading liquid NH3 received by
truck or rail and 250 psig storage tanks for a 30-day supply of NH3 are
included in this area. A steam-heated, shell-and-tube vaporizer is needed
for vaporization of NH3, before it is mixed with air. Air is supplied by a
blower to form a 5% NH3~in-air mixture for better dispersion in the flue gas.
The blower is sized based on a 5-inch water pressure drop. The NH3 injection
grid is primarily piping and the cost is incorporated into the area invest-
ment estimate shown in the summary capital investment estimates in the following
report section, Economic Evaluation and Comparison.
Gas Handling—
Two FD blowers are provided between the boiler economizer and the
reactors to overcome the pressure drop of the reactor system. A pressure
drop of 17.1 inches water for the system is used as a basis. Since these
fans are booster fans for the NOX removal system, the total investment and
revenue requirements are included in the FGT costs. Other fan requirements
for boiler and ESP units are not included.
Reactor Section—
Eight reactors are used, six of which are on stream while two are in the
regeneration stage. The reactors are parallel-passage design with the cata-
lyst contained in unit cells. The flue gas is passed over the catalyst
surface rather than through it. Reactor operating sequences are controlled
by 16 large flue gas valves, a hydraulic valve switching system and a
sequence controller. A steam superheater between the reactors and air
heater heats steam for purging during catalyst regeneration. Also included
in this section is the incremental investment required for a larger air heater
to recover portions of the claimed heat credits.
Flow Smoothing—
The off-gas from regeneration requires treatment for heat recovery,
water condensation to produce an S02~rich stream, and dampening of cyclic
flow. A waste heat boiler, steam disengaging drum, and steam blowdown drum
are used in the transfer of heat from the regeneration off-gas to boiler feed-
water. Quench towers and strippers are used in water condensation and the
production of a concentrated S02 gas stream. A gas compressor and gas
holder are used in dampening the flow of the S02-rich gas to the sulfuric
acid manufacturing plant.
Steam-Naphtha Reformer—
Hydrogen gas for regeneration of the reactors is manufactured in a
steam-naphtha reformer producing about 10 Msft3/day of 95% purity hydrogen.
The naphtha requirement is 1,500 gal/hr (assuming 80% for feed and 20% for
fuel). The electricity required is 250 kW and 12.4 klb/hr of steam is
produced for other uses.
126
-------
Sulfuric Acid Plant—
There are several possible uses of the S02~rich off-gas from the regen-
eration process, i.e., conversion to liquefied SC>2, elemental sulfur, or
H2S04- In this study, on the process developer recommendation, H2S04 is
produced. A contact unit producing 427 tons/day of 100% I^SO^ as 98% acid
is used. The electricity required is 2,050 kW and the byproduct steam
available for other uses is 32 Mlb/hr. The flue gas from the sulfuric acid
plant meets EPA standards. Savings could be realized if a less expensive
acid plant were used in which the flue gas from the acid plant would be
combined with the boiler flue gas for S0x-N0x FGT.
LIMESTONE FGD
The limestone FGD system with four trains of absorbers is designed for
90% removal of S02- The system does not include an ESP or a prescrubber.
Each of the absorbers is a TCA type with a presaturator for cooling and
humidifying the flue gas. Each scrubber is equipped with a chevron-type
entrainment separator at the scrubber outlet. Stack gas reheat to 175°F
for plume buoyancy is provided by indirect steam reheat of the cleaned gas.
Waste sludge is disposed of by ponding as a 15 weight percent solids slurry.
The settled sludge contains 60% free water; the excess water is returned to
the scrubber system.
A complete process description, flowsheet, material balance, and equip-
ment list for the limestone FGD process used in this study may be found in
the following EPA report, Definitive SOx Control Process Evaluations:
Limestone, Double Alkali^ and^ Citrate FGD Processes, by Tomlinson, S. V.,
et al (17).
ESP
Four ESP's of varying efficiencies are included in this study to allow
overall SOx-NOx-PM control system comparability. Four 99.5% efficient cold
ESP's are used with the Hitachi Zosen and UOP SFGT-N processes. Each ESP
has about 117,000 ft2 of plate area and handles about 391,000 aft3/min of
flue gas at 300°F. The estimated total equipment cost in 1979 dollars is
$6,141,000.
Four larger 99.5% efficient cold ESP's are used with the UOP SFGT-SN
process since this process also removes some 803, thus increasing the
difficulty for fly ash removal. Each ESP has about 180,000 ft2 of plate
area and handles about 390,000 aft3/min of flue gas at 30QOF. The estimated
total equipment cost in 1979 dollars is $7,934,000.
Four 98% efficient hot ESP's are used with the Kurabo Knorca process.
The remaining 1.5% PM is removed downstream at the NOx reactor. Each ESP
has 129,000 ft2 of plate area and treats about 537,000 aft3/min of flue gas
at 705°F. The estimated total equipment cost in 1979 dollars is $6,924,000.
127
-------
One 94% efficient cold ESP is used with the Moretana Calcium process
while the remaining PM is removed in the downstream prescrubber. This ESP
has 232,000 ft2 of collecting area and handles about 1,544,000 aft3/min of
flue gas at 300°F. The estimated total equipment cost in 1979 dollars is
$4,429,000.
128
-------
ECONOMIC EVALUATION AND COMPARISON
Based on the definitive power plant, process design, and economic
premises outlined earlier in this report and on the specific process equip-
ment for each process described in the previous section, total capital
investment and total annual revenue requirements were prepared for the
economic evaluation and comparison of the seven alternative NOX FGT systems.
It should be emphasized once again that in some cases, particularly the dry
N0x-only processes, additional costs for FGD and ESP must be added to the '
capital investment and annual revenue requirement for N0x-only removal to
provide an equal comparison between the NOX FGT alternatives.
PROCEDURE FOR ESTIMATING THE TOTAL CAPITAL INVESTMENTS
From material balance calculations and information supplied by the
process vendors, the various pieces of equipment for each process were sized
and costed. The process equipment costs were obtained primarily by using
in-house cost data but vendor quotes were used for unique or unusual equip-
ment. This cost for each piece of equipment reflected the midwestern power
plant location and was escalated to mid-1979 dollars by assuming an escala-
tion rate of 8%/yr. Tables containing the equipment list, equipment
description, and the mid-1979 equipment cost (material and labor) for each
process are included in the Systems Estimated section.
Because of the preliminary nature of this study, the various installa-
tion expenses such as piping, foundations, etc., were estimated as a
percentage of equipment costs based on both accepted ranges from published
data and previous TVA experience. The subtotal direct investment is the
sum of the process equipment costs and installation expenses. The final
component of the total direct investment is the services and miscellaneous
cost which were calculated as percentages of the subtotal direct investment.
The sum of the subtotal direct investment plus the services and miscellaneous
cost is the total direct investment.
The total indirect investment (the sum of various indirect investment,
such as engineering design and supervision, A&K contractor, construction
expense, and contractor fees) was calculated based on the equipment costs
or the total direct investment using the formulas described in the Economic
Premises section. A contingency calculated as 20% of the sum of the total
direct and total indirect investment was also included. The sum of the
total direct investment, the total indirect investment, and the contingency
is the total fixed investment.
129
-------
Other capital charges consist of an allowance for startup and modifica-
tion, interest during construction, land costs, working capital, and royalty
fee. The allowance for startup and modification and interest during con-
struction were calculated as a percentage of the total fixed investment.
The total depreciable investment is the sum of the total fixed investment,
the previously calculated allowance for startup and modification, and interest
during construction. The sum of total depreciable investment, land costs,
working capital, and royalty fee is the total capital investment.
RESULTS
The total capital investment for each of the seven NOX FGT alternatives
calculated for a base case, new, 500-MW coal-fired power plant is listed on
the following pages.
Asahi S0x-N0x Process
Total Capital Investment —
The total capital investment for the base case application for the Asahi
process was estimated at $104. 9M ($233/kW) in mid-1979 dollars. This total
cost can be broken down into the various investment cost categories shown in
Table 39. The direct investment accounts for 54.1% of the estimated total
capital investment for the Asahi process. The indirect investments make up
24.5% of the required investment and the various other capital charges make
up the remaining 21.4%.
The direct investment can be further broken down into costs by processing
area also shown in Table 39. The cracking, SOX-NOX absorption, and crystalliza-
tion areas alone account for about 62% of the total direct investment. The
high costs of these areas are due to the use of a liquid scrubbing system
and the resulting necessity of investment for large reactors, evaporators,
crystallizers, a dryer, rotary kiln crackers, and screw decanter separators.
Although the absorption section consists of four trains with only four pieces
of equipment in each train, the direct investment for this section is nearly
$11. 5M, reflecting the difficulty in absorbing low concentrations of NOx into
aqueous solution. This high investment cost for absorbers must be, however,
compared with an expensive generating system to produce a gas-phase oxidant
in the oxidation processes.
Major indirect investments were the construction expense at $7.1M,
engineering design and supervision at about $2.2M, and contractors fees at
$2.1M. A&E contractors expense was estimated at $0.5M. A 20% contingency
of $13. 8M Was assessed.
The remaining $22. 4M of the total capital investment was made up of
other capital charges, such as allowance for startup and modification,
interest during construction, land, working capital, and royalty fee. Ten
percent of the total fixed investment ($8.3M) was assessed as an allowance
130
-------
TABLE 39. ASAHI SOX-NOX PROCESS
SUMMARY OF ESTIMATED TOTAL CAPITAL INVESTMENT3
(500-MW new coal-fired power unit,
3.5% S in coal, 90% NOX removal, 99% S02 removal)
"/.. ill total
1 nvost mi'nb , $ _d^i rcLrj i iwi'st mi-iu^
Direct Investment*1
Materials handling (hoppers, conveyors, and dust collector) 1,013,000
Feed preparation (crusher, ball mill, tanks, agitators, and pumps) 1,288,000
Gas handling (booster fans and gas ducts and dampers from absorbers
to stack) 5,105,000
Particulate control (common feed plenum and gas ducts and dampers
from plenum to prescrubber, four prescrubbers, thickener, tanks,
agitators, and pumps) 2,617,000
SO -NO absorption (four absorbers, tanks, agitators, and pumps) 11,407,000
Crystallization (evaporators, crystallizers, refrigeration units,
screw decanters, tanks, agitators, and pumps) 11,395,000
Nitrogen treatment (screw decanter, cracker, filter, conveyors,
tanks, agitators, and pumps) 1,168,000
Cracking (dryer, crackers, furnace, blowers, conveyors, tanks,
and pumps) 12,328,000
Gypsum production (thickener, screw decanter, conveyor, tanks,
agitators, and pumps) 2,385,000
Absorbent regeneration (stripper, thickeners, conveyors, tanks,
agitators, and pumps) 3,684,000
Raw material storage (hoppers, conveyors, tanks, agitator,
and pumps) 850,000
Solids disposal . 46.000
Subtotal 53.286,000
Services, utilities, and miscellaneous 3,197,000
Trucks and earthmoving equipment 328,000
Total direct investment 56,811.000
I.H
2.3
9.0
20. 1
20.0
2.0
Jl. 7
4.2
6.5
.1.5
0.1
93.8
5.6
0.6
100.0
Indirect Investment
Engineering design and supervision
Architect and engineering contractor
Construction expense
Contractor fees
Total indirect investment
Contingency
Total fixed investment
2,181,000
545.000
7,147,000
2.068.000
11,941,000
13,750,000
82,502,000
3.8
1.0
12.6
3.6
21.0
24.2
145.2
Other Capital Charges
Allowance for startup and modification
Interest during construction
Total depreciable investment
Land
Working capital
Royalty fee
Total capital investment
8,250,000
9.900.000
100.652,000
254,000
3,703,000
300.000
104,909,000
$/aft3/mln $/kW
48.9 233
14.5
17.5
177.2
0.5
6.5
0.5
184.7
Basis
Midwest plant location represents project beginning mid-1977, ending mid-1980. Average cost
basis for scaling, mid-1979.
Stack gas reheat to 175°F by indirect steam reheat.
Minimum in-process storage; only pumps are spared.
Disposal pond located 1 mile from power plant.
Investment requirements for fly ash removal and disposal excluded; Ff.D process investment
estimate begins with common feed plenum downstream of the ESP.
Construction labor shortages with accompanying overtime pay incentive not considered.
Each area direct investment includes total equipment costs (given in the equipment lists) plus
installation labor and material costs for electrical, piping, ductwork, foundations, structural,
instrumentation, Insulation, and site preparation.
131
-------
for plant startup and modifications. Interest during construction was
estimated as 12.0% of the total fixed investment, or $9.9M. The total cost.
for land was estimated at about $0.25M and working capital at $3.7M. A
royalty fee of $0.3M was also added.
IHI SOX-NOX Process
Total Capital Investment—
The estimated total capital investment for the IHI SOx-NOx process was
$203.6M ($482/kW) in mid-1979 dollars as shown in Table 40. The direct
investment accounts for 57.6% of the total capital investment. The indirect
investments make up 22.5% of the total investment and the various other
capital charges make up the remaining 19.9%.
The direct investment for the IHI process can be further subdivided
into the various processing areas as shown in Table 40. The direct invest-
ment for the ozone generating system dwarfs the investments required for
the other processing areas, contributing 66%, about two-thirds of the total
direct investment and 38.1% of the total capital investment. Other capital-
intensive processing areas include: particulate control, $10.OM; SOX-NOX
absorption, $7.2M; and nitrate treatment, $5.2M.
Major indirect investments were the construction expense, $13.1M and
contractors fees at $3.6M. Engineering design and supervision and A&E .
contractor expense contributed significantly less at $1.5M and $0.4M
respectively. The remaining indirect investment, project contingency, was
$27.2M.
The remaining $40.5M of the total capital investment was made up of
other capital charges including: allowance for startup and modifications,
interest during construction, land, working capital, and royalty fee. The
allowance for startup and modifications and interest during construction
were assessed at 10% ($16.3M), and 12% C$19.6M), respectively, and contri-
buted most of the other capital charges. Based on information from the
process vendors on land requirements and from material balance calculations,
land and working capital vere estimated as $0.29M and $4.0M. The royalty
fee added was $0.3M.
Moretana Calcium SOx-NOx Process
Total Capital Investment—
The total capital investment of the Moretana Calcium process was $88.OM
($190/kW). This can be broken down into the direct investment, the indirect
investment, and other capital charges as shown in Table 41. The direct
investment at slightly more than $46.7M accounts for 53.1% of the total
capital investment. The indirect investments are 24.5% of the toal capital
investment, while other capital charges constitute the remaining 22.4%.
The direct investment for the Moretana Calcium process can be further
subdivided into the various processing areas shown in Table 41. The major
processing area, in terms of the direct investment involved, is the chlorine
dioxide generating system. A direct investment of nearly $16.8M, 19.1% of the
132
-------
TABLE AC. IHI SOX-NOX PROCESS
SUMMARY OF ESTIMATED TOTAL CAPITAL INVESTMENT3
(500-MW new coaJ-fired power unit,
3.5% S in coal, 90% NOX removal, 90% S02 removal)
% of total
Investment, $ direct Investment
Direct Investmentb
Material handling (hoppers, conveyors, and dust collector)
Feed preparation (crusher, ball mill, tanks, agitators, and pumps)
Gas handling (common plenum, ducts and dampers from plenum to pre-
scrubber, and booster fans)
Particulate control (prescrubber, thickeners, filter, tanks, agi-
tators, and pumps)
SO -NO absorption (absorber and pumps)
Stack gas reheat (single indirect steam reheater)
Gypsum production (oxidation tower, thickener, centrifuge, tanks,
agitators, and pumps)
Nitrate treatment (evaporators, thickener, decomposer, blowers,
conveyors, furnace, tanks, agitators, and pumps)
Raw material storage (silo, conveyors, tanks, agitators, and pumps)
Solids disposal
Ozone generating system (one unit complete)
Subtotal
Services, utilities, and miscellaneous
Trucks and earthmoving equipment
Total direct investment
1,013,000
1,153,000
4,371,000
9,952,000
7,239,000
1,016,000
2,546,000
5,153,000
328,000
50,000
77,562.000
110,383,000
6,623,000
36A.OOP
117,370,000
0.8
1.0
3.7
8.5
6.2
0.9
2.2
4.4
0.3
66.1
94.1
5.6
0.3
100.0
Indirect Investment
Engineering design and supervision
Architect and engineering contractor
Construction expense
Contractor fees
Total indirect Investment
Contingency
Total fixed investment
1,540,000
385,000
13,052,000
3,590,000
18,567,000
27.187.000
163,124,000
1.3
0.3
11.1
3.1
15.8
23.2
139.0
Other Capital Charges
Allowance for startup and modification
Interest during construction
Total depreciable investment
Land
Working capital
Royalty fee
Total capital Investment;
16,312,000
19,575.000
199,011,000
285,000
4,026,000
300.000
203,622,000
$/aft3/min $/kW
94.9 482
13.9
16.7
169.6
0.2
3.4
0.3
173.5
Basle
Midwest plant location represents project beginning mid-1977, ending mid-1980. Average cost
basis for scaling, mid-1979.
Stack gas reheat to 175°F by Indirect steam reheat.
Minimum in-process storage; only pumps are spared.
Disposal pond located 1 mile from power plant.
Investment requirements for fly ash removal and disposal excluded; FGD process investment
estimate begins with common feed plenum downstream of the ESP.
Construction labor shortages with accompanying overtime pay Incentive not considered.
bach area direct investment includes total equipment costs (given in the process equipment list)
plus installation labor and material costs for electrical, piping, ductwork, foundations,
structural, instrumentation, insulation, and site preparation.
133
-------
TABLE 41. MORETANA CALCIUM PROCESS
SUMMARY OF ESTIMATED TOTAL CAPITAL INVESTMENT3
(500-MW new coal-fired power unit,
3.5% S in coal, 90% NOX removal, 95% S02 removal)
% of total
Investment, $ direct investment
Direct Investment^
Material handling (hoppers, conveyors, and dust collector)
Feed preparation (crusher, ball mill, tanks, agitators, and pumps)
Gas handling (common plenum, ducts and dampers from plenum to
prescrubber, and booster fans)
Particulate control (two prescrubbers, tanks, agitators, and
pumps)
SOx-NOx absorption (two absorbers, clarifier, tanks, agitators,
and pumps)
Stack gas reheat (two indirect steam reheaters)
Gypsum production (centrifuge, conveyor, tanks, agitators,
and pumps)
Nitrogen treatment (centrifuge, evaporators, conveyor, tanks,
agitators, and pumps)
Raw material storage (silos, hoppers, and conveyors)
Solids disposal (sludge pond, tanks, agitators, and pumps)
C102 generation (two skid-mounted units complete)
Subtotal
Services, utilities, and miscellaneous
Total direct investment
1,111,000
1,461,000
4,246,000
1,591,000
4,640,000
1,359,000
2,732,000
2,618,000
918,000
6,563,000
16.80_9,QOQ
44,048,000
2.643.000
46.691,000
2.4
3.1
9.1
3.4
9.9
2.9
5.8
5.6
2.0
14.1
36.0
94.3
5.7
100.0
Indirect Investment
Engineering design and supervision
Architect and engineering contractor
Construction expense
Contractor fees
Total indirect investment
Contingency
Total fixed investment
1,955,000
440,000
5,975,000
1.782.000
10,161,000
11,370,000
68,222,000
4.2
1.0
12.8
3.8
21.8
24.3
146.1
Other Capital Charges
Allowance for startup and modification
Interest during construction
Total depreciable investment
Land
Working capital
Royalty fee
Total capital investment
6,166,000
8,187,000
82,5/5,000
910,000
4,182,000
300,000
87,967,000
$/aft3/min $/kW
41.0 190
13.2
17.5
176.8
2.0
9.0
0.6
188.4
a. Basis
Midwest plant location represents project beginning mid-1977, ending mid-1980. Average cost
basis for scaling, mid-1979.
Stack gas reheat to 175°F by indirect steam reheat.
Minimum in-process storage; only pumps are spared.
Disposal pond located 1 mile from power plant.
Investment requirements for fly ash removal and disposal excluded; FGD process investment
estimate begins with common feed plenum downstream of the ESP.
Construction labor shortages with accompanying overtime pay incentive not considered.
b. Each area direct investment includes total equipment costs (given in the process equipment list)
plus'installation labor and material costs for electrical, piping, ductwork, foundations,
structural, instrumentation, insulation, and site preparation.
134
-------
total capital investment, is required. Other capital-intensive processing
areas are SOx-NOx absorption, $4.6M and solids disposal, $6.6M.
Major indirect investments are construction expense, $6.0M; engineering
design and supervision, $2.0M; and contractors fees, $1.8M. A&E contractor
expense, $0.50M, was significantly less. The project contingency was $11.4M.
The allowance for startup and modifications and interest during construc-
tion were assessed at 10% ($6.2M), and 12% ($8.2M), respectively, of the total
fixed investment. These two capital charges together contributed nearly
16.3% of the total capital investment and made up most of the other capital
charges. Based on information from the process vendor on land requirements
and from material balance calculations, land and working capital were
estimated at $0.9M and $4.2M. A $0.3M royalty fee was added.
Hitachi Zosen NOx-Only Process
The total capital investment for the Hitachi Zosen NOjj-only process was
estimated to be $23.3M ($48.2/kW). Table 42 shows a breakdown of this total
capital investment. The direct investment composes 53.4% of the estimated
total capital investment. The indirect investment constitutes 26.2% of the
total capital investment. The other capital charges compose the remaining
20.4%.
The direct investment comprises the three major processing area costs
shown in Table 42. The cost for the reactor section accounts for 83.7% of
the total direct investment. This section, as for most of the other dry NOX
removal processes, is the heart of the process.
The major indirect investments were the construction expense, $2.0M
and contractor fees, $0.7M. Both were based on the total direct investment.
The engineering design and supervision at $0.3M and the A&E contractor
expense at $0.07M were based on the number of pieces of equipment. A contin-
gency of $3.1M was also assessed.
Under other capital charges, allowance for startup and modifications
was $1.9M which is 10% of the total fixed investment. Interest during
construction was estimated as 12.0% of the total fixed investment of $2.2M.
Land cost was estimated at $5,000 while working capital was estimated at
$0.4M. A $0.3M royalty fee, as stated by the vendor, was added.
Kurabo Knorca NOy-Only Process
The total capital investment of the Kurabo Knorca process was estimated
to be $21.2M ($43.9/kW). This total capital investment is broken down into
direct investment, indirect investment, contingency, and other capital
charges as shown in Table 43. The direct investment composes 51.4% of the
total capital investment, or slightly less than $10.9M. The indirect
investment and contingency account for 28.2% of the total capital investment,
or about $6.0M. The other capital charges make up the remaining 20.4%.
135
-------
TABLE 42. HITACHI ZOSEN PROCESS
SUMMARY OF ESTIMATED CAPITAL INVESTMENT3
(500-MW new coal-fired power unit, 3.5% S in coal;
90% NOx removal)
Direct Investment^
NH3 storage and injection (compressors, tanks,
vaporizers, and blowers)
Reactor section (reactors, catalyst fill)
Gas handling (flue gas fans)
Subtotal
Services, utilities, and miscellaneous
Total direct investment
Investment, $
% of
total direct
investment
688,000
10,419,000
635.000
11,742,000
705,000
12,447,000
5.5
83.7
5.1
94.3
5.7
100.0
Indirect Investment
Engineering design and supervision
Architect and engineering contractor
Construction expense
Contractor fees
Total indirect investment
Contingency
Total fixed investment
274,000
69,000
2,027,000
652.000
3,022,000
3.094.000
18,563,000
2.2
0.6
16.3
5.2
24.3
24.8
149.1
Other Capital Charges
Allowance for startup and modifications
Interest during construction
Total depreciable investment
Land
Working capital
Royalty fee
Total capital investment
1,856,000
2,228.000
22,647,000
5,000
355,000
300.000
23,307,000
$/aft3/min $/kW
10.9 48.2
14.9
17.9
181.9
0.0
2.9
2.4
187.2
a. Basis
Midwest plant location. represents project beginning mid-1977, ending
mid-1980. Average basis for scaling, mid-1979.
Investment requirements for fly ash disposal excluded.
Construction labor shortages with accompanying overtime pay incentive
not considered.
b. Each area direct investment includes total equipment costs (given in the
process equipment list) plus installation labor and material costs for
electrical, piping, ductwork, foundations, structural, instrumentation,
insulation, and site preparation.
136
-------
TABLE 43. KURABO KNORCA PROCESS
SUMMARY OF ESTIMATED CAPITAL INVESTMENT3
(500-MW new coal-fired power unit, 3.5% S in coal;
90% NOx removal)
% of
total direct
Investment, $ investment
Direct Investment"
NH3 storage and injection (compressors, tanks,
and vaporizers) 691,000
Reactor section (static mixer, reactor, feeder,
conveyors, hopper, elevator, and catalyst fill) 8,898,000
Gas handling (flue gas fans) 635,000
Waste disposal (filters and blowers) 58,OOP
Subtotal 10,282,000
Services, utilities, and miscellaneous 617.000
Total direct investment 10,899,000
6.3
81.7
5.8
0.5
94.3
5.7
100.0
Indirect Investment
Engineering design and supervision
Architect and engineering contractor
Construction expense
Contractor fees
Total indirect investment
Contingency
Total fixed investment
599,000
150,000
1,815,000
590.000
3,15A,000
2,811,000
16,864,000
5.5
1.4
16.6
5.4
28.9
25.8
154.7
Other Capital Charges
Allowance for startup and modifications
Interest during construction
Total depreciable investment
Land
Working capital
Royalty fee
Total capital investment
1,686,000 15.5
2,024,000 18.6
20,574,000 188.8
5,000 0.0
311,000 2.9
300,000 2.7
21,190,000 194.4
$/aft3/min $/kW
9.87 43.9
a. Basis
Midwest plant location represents project beginning mid-1977, ending
mid-1980. Average basis for scaling, mid-1979.
Investment requirements for fly ash disposal excluded.
Construction labor shortages with accompanying overtime pay incentive
not considered.
b. Each area direct investment includes total equipment costs (given in the
process equipment list) plus installation labor and material costs for
electrical, piping, ductwork, foundations, structural, instrumentation,
insulation, and site preparation.
137
-------
The direct investment for the Kurabo Knorca process may be subdivided
as shown in Table 43. The reactor section, which contains the reactors, cata-
lyst, static gas mixer, catalyst feeders and hoppers, conveyors, and bucket
elevators, is the most costly area at almost $8.9M. This represents 81.7%
of the total direct investment.
The construction expense, the major indirect investment of $1.8M, and
contractor fees of $0.6M were based on the total direct investment. The
engineering design and supervision, estimated at about $0.6M, and A&E
contractor expense, estimated at $0.2M, were based on the number of pieces
of equipment. The contingency was $2.8M.
The other capital charges totaled about $4.3M. Included were allowance
for startup and modification, $1.7M, and interest during construction, $2.0M,
which were estimated at 10% and 12%, respectively, of the total fixed invest-
ment. Land and working capital were $5,000 and $0.3M respectively. A royalty
fee of $0.3M was also added.
UOP SFGT-N, NOx-Only Process
The total capital investment of the UOP SFGT-N process was estimated to
be about $18.4M ($38.1/kW). Table 44 shows the subdivision of this total
capital investment. The direct investment represents 53.4% of the total
capital investment, the indirect investment and contingency are 27.0% and
other capital charges are the remaining 19.6%.
The direct investment comprises the three major processing area costs
shown in Table 44. The cost for the reactor section accounts for 80.9% of
the total direct investment. This section contains the reactors, unit cells,
and initial catalyst fill.
The major indirect investments were the construction expense of about
$1.7M and contractor fees of $0.5M. The engineering design and supervision
of $0.2M and the A&E contractor expense of $0.06M were derived from the
number of pieces of equipment. The contingency was $2.5M.
Other capital charges include allowance for startup and modification,
interest during construction, land, and working capital. Allowance for
startup and modification was taken as 10% of the total fixed investment, or
about $1.5M. Interest during construction was estimated as 12% of the total
fixed investment, or $1.8M. Land and working capital were estimated to be
$5,000 and $0.3M respectively. The royalty fee was prorated and included
under the direct investment.
UOP SFGT-SN. SOy-NOx Process
The total capital investment for the UOP SFGT-SN process was estimated
to be $67.2M ($139/kW). Table 45 shows a breakdown of this total capital
investment. The direct investment and indirect investment plus contingency
account for 55.6% and 24.4%, respectively, of the total capital investment.
The remaining 20% comprises the other capital charges.
138
-------
TABLE 44. UOP SFGT-N, NOX-ONLY PROCESS
SUMMARY OF ESTIMATED CAPITAL INVESTMENT3
(500-MW new coal-fired power unit, 3.5% S in coal;
90% NOX removal)
% of
total direct
Investment, $ investment
Direct Investment**
NH3 storage and injection (compressors, tanks,
vaporizers, and blowers)
Reactor section (reactors, catalyst fill)
Gas handling (flue gas fans)
Subtotal
Services, utilities, and miscellaneous
Total direct investment
Indirect Investment
Engineering design and supervision
Architect and engineering contractor
Construction expense
Contractor fees
Total indirect investment
Contingency
Total fixed investment
Other Capital Charges
Allowance for startup and modifications
Interest during construction
Total depreciable investment
Land
Working capital
Total capital investment
688,000
7,964,000
635,000
9,287,000
557,000
9,844,000
236,000
59,000
668,000
546,000
2,509,000
2,471,000
14,824,000
1,482,000
1,779,000
18,085,000
5,000
333,000
18,423,000
$/aft3/min $/kW
7.0
80.9
6.4
94.3
5.7
100.0
2.4
0.6
16.9
5.6
25.5
25.1
150.6
15.0
18.1
183.7
0.0
3.4
187.1
a. Basis
Midwest plant location represents project beginning mid-1977, ending
mid-1980. Average basis for scaling, mid-1979.
Investment requirements for fly ash disposal excluded.
Construction labor shortages with accompanying overtime pay incentive
not considered.
b. Each area direct investment includes total equipment costs (given in the
process equipment list) plus installation labor and material costs for
electrical, piping, ductwork, foundations, structural, instrumentation,
insulation, and site preparation.
139
-------
TABLE 45. UOP SFGT-SN, SOX-NOX PROCESS
SUMMARY OF ESTIMATED CAPITAL INVESTMENT3
(500-MW new coal-fired power unit, 3.5% S in coal;
90% NCbc removal, 90% S02 removal)
Direct Investment"
NH3 storage and Injection (compressors, tanks, vaporizers, and
blowers)
H2S04 plant (contact unit)
Reactor section (reactors, catalyst fill, superheater, valves,
hydraulic valve switching system, unit cells)
Flow smoothing section (boiler, drums, quench towers, pumps,
filters, strippers, fans, and compressors)
Steam naphtha reformer (one reformer unit)
Gas handling (flue gas fans)
Subtotal
Services, utilities, and miscellaneous
Total direct investment
Investment, $
% of total
direct investment
688,000
7,172,000
19,110,000
2,545,000
4,105,000
1,589.000
35,209,000
2.113.000
37,322,000
1.8
19.2
51.2
6.8
11.0
4.3
94.3
5.7
100.0
Indirect Investment
Engineering design and supervision
Architect and engineering contractor
Construction expense
Contractor fees
Total indirect investment
Contingency
Total fixed investment
709,000
177,000
5,043,000
1.503.000
7,432,000
8.951.000
53,705,000
1.9
0.5
13.5
4.0
19.9
24.0
143.9
Other Capital Charges
Allowance for startup and modifications
Interest during construction
Total depreciable investment
Land
Working capital
Total capital investment
5,371,000 14.4
6.445,000 17.3
65,521,000 175.6
14,000 0.0
1,618.000 4.J
67,153,000 179.9
$/aft3/min $/UW
31.3 139
a. Basis
Midwest plant location represents project beginning mid-1977, ending mid-1980. Average
basis for scaling, mid-1979.
Investment requirements for fly ash disposal excluded.
Construction labor shortages with accompanying overtime pay incentive not considered.
b. Each area direct investment includes total equipment costs (given in the process equipment
list) plus installation labor and material costs for electrical, piping, ductwork, founda-
tions, structural, instrumentation, insulation, and site preparation.
140
-------
The direct investment is to be subdivided into the cost areas as shown
in Table 45. The reactor section, containing reactors, unit cells} initial
catalyst fill, steam superheater, hydraulic valve switching system and valves,
is the single largest direct investment factor, i.e., $19.1M or 51.2% of the
total direct investment. The sulfuric acid plant, steam naphtha reformer,
and flow smoothing sections are also capital-intensive process areas.
Major indirect investments were the construction expense at about $5.0M
and contractor fees at $1.5M. Engineering design and supervision and A&E
contractor expenses were about $0.7M and $0.2M respectively. Contingency
was estimated to be $9.0M.
The other capital charges were approximately $5.4M for allowance for
startup and modifications, $6.4M for interest during construction, $14k for
land, and $1.6M for working capital. The royalty fee was prorated and included
with the direct investment.
Limestone FGD
The total capital investment of the limestone FGD process with 90% S02
removal, which is used in addition to each of the dry N0x-only FGT processes
for overall system comparability with SOX-NOX processes, was estimated to be
$50.4M (17).
ESP
The total capital investment of the various ESP's was estimated using
the same procedures as previously given for the NOX processes. These
investments are summarized below.
1. The total capital investment, in 1979 dollars, for four 99.5%
efficient cold ESP's, to be used with the Hitachi Zosen, and UOP
SFGT-N processes, was estimated to be $10,830,000. This was based
on a total equipment cost of $6,141,000.
2. The total capital investment, in 1979 dollars, for four 99.5%
efficient cold ESP's, to be used with the UOP SFGT-SN process,
was estimated to be $14,625,000. This was based on a total
equipment cost of $7,934,000.
3. The total capital investment, in 1979 dollars, for four 98% efficient
hot ESP's, to be used with the Moretana Calcium process, was estimated
to be $12,146,000. This was based on a total equipment cost of
$6,924,000.
4. The total capital investment, in 1979 dollars, for one 94% efficient
cold ESP, to be used with the Moretana Calcium process, was estimated
to be $7,213,000. This was based on a total equipment cost of $4,429,000.
PROCEDURE FOR ESTIMATING THE AVERAGE ANNUAL REVENUE REQUIREMENTS
The annual revenue requirements were calculated based on a midwestern
power plant location and mid-1980 costs. These average annual revenue
141
-------
requirements are divided into the direct costs, which consist of raw
materials costs and conversion costs and indirect costs, which consist of
capital charges and overheads.
The raw material costs were calculated by determining the current (1978)
delivered price of these raw materials in the Chicago area and assuming an
escalation rate of 8%/yr for the delivered raw material. For the conversion
costs, the utility and labor costs were projected using internal TVA informa-
tion. Operating labor and supervision were based on both process vendor
information and previous TVA experience. The maintenance cost was calculated
as a percentage of the total direct investment based on previous TVA
experience.
The capital charges portion of the indirect costs consists of:
(1) depreciation, interim replacements, and insurance, which are based on
the total depreciable investment; and (2) the average costs of capital and
taxes, which are based on the total capital investment. The overhead charges
consists of: (1) plant overhead, which is based on a percentage of the cost of
operating labor and supervision, maintenance, and analyses; and (2) adminis-
trative overheads, which are based on a percentage of operating labor and
supervision. The methods by which these capital charges are determined were
described previously in the Economic Premises section.
The annual revenue requirements are the sum of the direct and the
indirect costs. Equivalent unit revenue requirements in mills/kWh can then
be obtained by dividing by the on-stream time of 7000 hr/yr and the plant
net MW rating.
RESULTS
Both the annual revenue requirements and the equivalent unit revenue
requirements were calculated for each of the seven NOX FGT alternatives
based on the design and economic premises of this study. A detailed
breakdown of these revenue requirements for each process is provided on
the following pages.
Asahi SOx-NOx Process
The total annual revenue requirement for the Asahi process is
approximately $39.8M as shown in Table 46. This corresponds to a unit
revenue requirement of 12.63 mills/kWh. Direct costs make up $22.OM
or slightly more than 55% of the total. Indirect costs, primarily capital
charges, account for the remaining 45.0%. The various capital charges of
$15.1M represent 37.8% of the total annual revenue requirements. Although
limestone is the largest in quantity, the major annual raw material costs
are for sodium carbonate and EDTA, $3.55M and $1.66M respectively. Major
conversion costs include fuel oil for cracking and flue gas reheat at $4.73M,
electricity (primarily for the flue gas blowers) at $4.3M, and plant
maintenance at $4.0M.
142
-------
TABLE 46. ASAHI SOX-NOX PROCESS
SUMMARY OF ANNUAL REVENUE REQUIREMENTS - REGULATED UTILITY ECONOMICS3
(500-MW new coal-fired power unit, 3.5% S in coal;
onsite solids disposal, 90% NOX removal, 99% S02 removal)
Annual
quantity
Direct Costs
Raw materials
Limestone 113,000 tons
FeSO.'7H20 840 tons
Na fiDTA 2,240 tons
Na,CO, 34,500 tons
K,SO/ 1,100 tons
H,S07 2,200 tons
£. 4
Total raw materials cost
Conversion costs
Operating labor and supervision
Plant 46,428 man-hr
Solids disposal equipment 29,120 man-hr
Utilities
Fuel oil 10,500,000 gal
Steam 663,000 MBtu
Process water 590,700 kgal
Electricity 147,000,000 kWh
Landfill operation
Trucks (fuel and maintenance) 198,200 tons
Earthmoving equipment (fuel and
maintenance) 198,200 tons
Maintenance
Labor and material
Analyses 8,760 man-hr
Total conversion costs
Total direct costs
Indirect Costs
Capital charges
Depreciation, interim replacements, and
insurance at 6.0% of total depreciable
Investment
Average cost of capital and taxes at 8.6%
of total capital investment
Overheads
Plant, 50% of conversion costs less utilities
Administrative, 10% of operating labor
Total indirect costs
Total annual revenue requirements
Equivalent unit revenue requirements
Total % of
Unit annual annual revenue
cost, $ cost, $ requirement
7.40/ton
86.50/ton
743.00/ton 1,
103.00/ton 3,
144.00/ton
54.00/ton
6,
12.50/man-hr
17. 00 /man-hr
0.45/gal 4,
2.00/MBtu 1,
0.12/kgal
0.029/kWh 4,
0.06/ton
0.16/ton
836,200
72,700
664,300
553,500
158,400
118,800
403,900
580,400
495,000
725,000
326,000
70,900
263,000
11,900
31,700
3,976,800
17. 00 /man-hr
15,
22,
6,
9,
2,
17,
39,
Mills /kWh
12.63
148,900
629,600
033,500
039.100
022,200
622,400
107,500
791,200
824,700
$/ton coal
burned
26.55
2.10
0.18
4.18
8.92
0.40
0.30
16.08
1.46
1.24
11.86
3.33
0.18
10.70
0.03
0.08
9.99
0.37
39.24
55.32
15.16
22.66
6.59
0.27
44.68
100.00
$/MBtu heat
input
1.26
Basis
Midwest plant location, 1980 revenue requirements.
Remaining life of power plant, 30 yr.
Power unit on-stream time, 7,000 hr/yr...
Coal burned, 1,500,100 tons/yr, 9,000 Btu/kWh.
Stack gas reheat to 175°F.
Investment and revenue requirements for disposal of fly ash excluded.
Total direct investment, $56,811,000; total depreciable investment, $100,652,000; and total
capital investment, $104,909,000.
143
-------
IHI SOX-NOX Process
The annual revenue requirement for the IHI process is slightly less
than $58.6M as shown in Table 47. This corresponds to a unit revenue
requirement of about 19.82 raills/kWh. Annual direct costs are $25.5M, about
43.5% of the total annual revenue requirements. Indirect costs, primarily
capital charges, account for the remaining 56.5%.
In contrast to the other wet,simultaneous SOX-NOX processes, the raw
material costs for the IHI process are $1.8M, a relatively insignificant
3.1% of the annual revenue requirements. Although large quantities of some
raw materials are used, they are relatively inexpensive materials such as
limestone. The major direct costs are conversion costs, primarily electri-
city although maintenance and steam consumption also add significant amounts.
Most of this electrical charge of $13.1M is consumed in generation of ozone.
The major annual costs associated with the IHI process are the two capital
charges (depreciation, etc., and average cost of capital) which together
account for more than 50% of the total annual revenue requirement.
Moretana Calcium SOx-NOx Process
The total annual revenue requirement for the Moretana Calcium process
is about $38.1M as shown in Table 48. This corresponds to an average unit
revenue requirement of 11.73 mills/kWh. Annual direct costs for raw
materials, labor, utilities, and maintenance are $23.6M or 62.1% of the.
total annual revenue requirements. Indirect costs, primarily for capital
charges, account for the remaining 37.9%.
Although the annual direct cost is similar to that of the other wet,
SOX-NOX processes, the raw material costs of $12.9M is significantly higher
than the other processes evaluated. Limestone, the primary raw material in
terms of quantity, is relatively cheap, but the Moretana Calcium process
also consumes several expensive raw materials. The annual cost of hydrogen
chloride (HC1) (used to generate C102) and several proprietary additives
contribute $4.3M and $4.1M, respectively, to the annual revenue requirements.
The conversion costs, which are lower for the Moretana Calcium process than
for the other wet processes, are about $10.7M, including electrical and
maintenance costs of $5.5M and $3.0M respectively. The capital charges of
nearly $12.5M represent 32.9% of the total annual revenue requirements.
Hitachi Zosen N0x-0nly Process
The total annual revenue requirement for the Hitachi Zosen process is
approximately $12.2M as shown in Table 49. This is equivalent to a unit
revenue requirement of about 3.61 mills/kWh. The direct costs, primarily
raw material costs, constitute about $8.8M annually, or 71.8% of the total
annual revenue requirements. Indirect costs, composed of capital charges
and overheads, form 30% of the annual revenue requirement or about $3.7M.
A credit for spent catalyst disposal of $0.26M, or 2.1% of the annual
revenue requirement, is applied.
144
-------
TABLE 47. IHI SOX-NOX PROCESS
SUMMARY OF ANNUAL REVENUE REQUIREMENTS - REGULATED UTILITY ECONOMICS3
(500-MW new coal-fired power unit, 3.5% S in coal;
onsite solids disposal, 90% NOX removal, 90% S02 removal)
Annual
quantity
Unit
cost, $
Total
annual
cost, $
% of
annual revenue
requirement
Direct Costs
Raw materials
Limestone
Powdered limestone
NaCI
Cud
NaOH
Total raw materials cost
Conversion costs
Operating labor and supervision
Plant
Solids disposal equipment
Utilities
Fuel oil
Steam
Process water
Electricity
Landfill operation
Trucks (fuel and maintenance)
Earthmoving equipment (fuel and
maintenance)
Maintenance
Labor and material
Analyses
Total conversion costs
Total direct costs
114,400 tons
770 tons
11,800 tons
1,160 tons
77 tons
2,200 tons
7.00/ton
54.00/ton
32.00/ton
25.80/ton
2,105.00/ton
187.00/ton
800,800
41,600
377,600
29,900
162,100
411.400
1,823,400
1.37
0.07
0.65
0.05
0.28
0.70
3.12
37,668 man-hr
32,000 man-hr
1,102,000 gal
1,463,000 MBtu
853,700 kgal
451,500,000 kWh
235,800 tons
235,800 tons
6,570 man-hr
12.50/man-hr
1 7 . 00/man-hr
0.45/gal
2.00/MBtu
0.10/kgal
0.029/kWh
0.06/ton
0.16/ton
17. 00/man-hr
470,900
544,000
495,900
2,926,000
85,400
13,093,500
14,100
37,700
5,889,000
111,700
23,668,200
25,491,600
0.80
0.93
0.85
5.00
0.15
22.35
0.02
0.06
10.05
0.19
40.40
43.52
Indirect Costs
Capital charges
Depreciation, interim replacements, and
insurance at 6.0% of total depreciable
investment
Average cost of capital and taxes at 8.631
of total capital investment
Overheads
Plant, 502 of conversion costs less utilities
Administrative, 10% of operating labor
Total indirect costs
Total annual revenue requirements
11,940,700
17,511,500
3,533,700
101.500
33,087,400
58.57ST.OOO
20.39
29.89
6.03
0.17
56.48
100.00
Equivalent unit revenue requirements
$/ton coal S/MBtu heat
Mills/kWh burned input
19.82
39.05
1.86
a. Basis
Midwest plant location, 1980 revenue requirements.
Remaining life of power plant, 30 yr.
Power unit on-stream time, 7,000 hr/yr.
Coal burned, 1,500,100 tons/yr, 9,000 Btu/kWh.
Stack gas reheat to 175°F.
Investment and revenue requirements for disposal of fly ash excluded.
Total direct investment, $117,370,000; total depreciable Investment, $199,011,000; and total
capital investment, $203,622,000.
145
-------
TABLE 48. MORETANA CALCIUM PROCESS
SUMMARY OF ANNUAL REVENUE REQUIREMENTS - REGULATED UTILITY ECONOMICS3
(500-MW new coal-fired power unit, 3.5% S in coal;
onsite solids disposal, 90% NOx removal, 95% S02 removal)
Direct Costs
Raw materials
Limestone
Slaked lime
CuCl
HC1
Proprietary additives
Total raw materials cost
Conversion costs
Operating labor and supervision
Utilities
Steam
Process water
Electricity
Maintenance
Labor and material
Analyses
Total conversion costs
Total direct costs
Annual
quantity
Unit
cost, $
Total
annual
cost, $
% of
annual revenue
requirement
172,000 tons
14,000 tons
56 tons
52,900 tons
28,250 tons
2,380 tons
7.00/ton
60.00/ton
2,105.00/ton
81.00/ton
82.00/ton
46,428 man-hr 12.50/man-hr
736,100 MBtu
365,300 kgal
189,280,000 kWh
8,760 man-hr
1,204,000
840,000
117,900
4,284,900
2,316,500
4.116.200
12,879,500
580,400
2.00/MBtu
0.12/kgal
0.029/kWh
17. 00 /man-hr
1,472,200
43,800
5,489,100
3,005,900
148,900
10,740,300
23,619,800
3.16
2.21
0.31
11.26
6.09
10.81
33.84
1.52
3.87
0.11
14.42
7.90
0.39
28.21
62.05
Indirect Costs
Capital charges
Depreciation, interim replacements, and
insurance at 6.0% of total depreciable
investment
Average cost of capital and taxes at 8.6%
of total capital Investment
Overheads
Plant, 50% of conversion costs less utilities
Administrative, 10% of operating labor
Total indirect costs
Total annual revenue requirements
4,954,500
7,565,200
1,867,600
58.000
14,445,300
38,065,100
13.02
19.87
4.91
O.j.5
37.95
100.00
Equivalent unit revenue requirements
Mills/kWh
11.73
$/ton coal
burned
25.38
$/MBtu heat
input
1.21
Basis
Midwest plant location, 1980 revenue requirements.
Remaining life of power plant, 30 yr.
Power unit on-stream time, 7,000 hr/yr.
Coal burned, 1,500,100 tons/yr, 9,000 Btu/kWh.
Stack gas reheat to 175°F.
Investment and revenue requirements for disposal of fly ash excluded.
Total direct investment, $46,691,000; total depreciable investment, $82,575,000; and total
capital investment, $87,967,000.
146
-------
TABLE 49. HITACHI ZOSEN PROCESS
SUMMARY OF ANNUAL REVENUE REQUIREMENTS - REGULATED UTILITY ECONOMICS*
(500-MW new coal-fired power unit, 3.5% S in coal;
90% NOX removal)
Direct Costs
Raw materials
NH3
Catalyst
Total raw materials costs
Conversion costs
Operating labor and supervision
Utilities
Steam
Electricity
Heat credit
Maintenance
Labor and material
Analyses
Total conversion costs
Total direct costs
Annual quantity Unit cost, 5
6,395 short tons 150.00/short ton
8,760 man-hr 12.50/man-hr
8,096 MBtu 2.00/MBtu
15,796,000 kWh 0.029/kWh
100,310 MBtu 2.00/MBtu
2,920 man-hr 17.00/man-hr
% of
Total annual annual revenue
cost, $ requirement
959,300
6,880,000
7,839,300
109,500
16,200
458,100
(200,600)
497,900
49,600
930,700
8,770,000
7
56
64
0
0
3
(1
i,
0
7
71
.85
.34
.19
.90
.13
.75
.64)
.07
.Al
.62
.81
Indirect Costs
Capital charges
Depreciation, interim replacements, and
insurance at 6. OX of total depreciable
investment
Average cost of capital and taxes at 8.6%
of total capital investment
Overheads
Plant, 50% of conversion costs less utilities
Administrative, 10% of operating labor
Total indiiect costs
Spent cataylst disposal.
Total average annual revenue requirements
Equivalent unit revenue requirements for NO* removal
1,358,800
2,004,400
328,500
11,000
3,702,700
(260.000)
12,212,700
$/ton
Kllls/kWh coal burned
3.61 8.14
11.13
16.41
2.69
0.09
30.32
(2.13)
100.00
S/MBtu
heat input
0.39
Basis
Midwest plant location, 1980 revenue requirements.
Remaining life of power plant, 30 yr.
Power unit on-stream time, 7,000 hr/yr.
Coal burned, 1,500,100 tons/yr, 9,000 Btu/kWh.
Investment and revenue requirement for disposal of fly ash excluded.
Total direct investment, $12,447,000! total depreciable investment, $22,647,000; and total capital invest-
ment, 523,307,000.
147
-------
As in all the dry, N0x-only processes, NH3 is the raw material used in
the largest quantity. However, the major raw material cost is for catalyst,
i.e., about $6.9M annually. The catalyst cost represents approximately 56.3%
of the annual revenue requirements and £s the largest single item in the
annual revenue requirements. The major conversion costs are maintenance at
almost $0.50M annually and electricity (primarily for the flue gas blowers)
at $0.46M. A heat credit valued at $0.2M is applied based on the heat of
reaction from the reduction of NOX with NH3 and the heat released from the
combination of excess NH3 with oxygen, less the preheat required for the NH3
and air from the NH3 injection system. The net heat released was considered
to be recovered at the air heater.
The capital charges account for about 27.5% of the total annual revenue
requirements, i.e., about $3.4M. Overheads make up 2.8% or $0.3M of the
annual revenue requirements. The $0.26M credit for spent catalyst disposal
is based on the scrap metal value of the catalyst base support.
Kurabo Knorca NQx-Only Process
The total annual revenue requirement for the Kurabo Knorca process is
approximately $9.3M as shown in Table 50. This is equivalent to a unit
revenue requirement of 2.77 mills/kWh. The direct costs, composed of raw
material and conversion costs, constitute about $6.0M annually or 63.9% of
the total annual revenue requirements. Indirect costs, composed of capital
charges and overheads, are 36.0% of the annual revenue requirements or about
$3.AM. The spent catalyst disposal cost of $6k is less than 0.1% of the
annual revenue requirement.
NH3 is the raw material used in the largest quantity. However, the
major raw material cost is for catalyst, i.e., about $4.3M annually. The
catalyst cost represents approximately 46.2% of the annual revenue require-
ments. The major conversion costs are plant maintenance at $0.4M annually
and electricity (primarily for flue gas blowers) at about $0.5M. A heat
credit valued at $0.2M is applied based on the heat of reaction from the
reduction of NOx with NH3 plus the heat released from the combination of
excess NH3 with oxygen, less the preheat required for the NH3 and air from
the NH3 injection system. The net heat released was considered to be
recovered in the air heater.
The capital charges are a major annual expense, accounting for $3.1M,
or about 32.7%. of the total annual revenue requirements. Overheads make up
$0.3M or 3.3% of the annual revenue requirements. The spent catalyst
disposal cost, based on disposal cost for Ti02, the catalyst carrier,*is
estimated at $6k.
UQP SFGT-N. NOy-Only Process
The total annual revenue requirement for the UOP SFGT-N process is
approximately $7.2M as shown in Table 51. This is equivalent to a unit
revenue requirement of 2.13 mills/kWh. The direct costs, comprising raw
material and conversion costs, are about $4.2M, 58.8% of the total annual
revenue requirements. Indirect costs, comprising capital charges and
148
-------
TABLE 50. KURABO KNORCA PROCESS
SUMMARY OF ANNUAL REVENUE REQUIREMENTS - REGULATED UTILITY ECONOMICS3
(500-MW new coal-fired power unit, 3.5% S in coal;
90% NOX removal)
Direct Costs
Raw materials
NH3
Catalyst
Total raw materials costs
Conversion costs
Operating labor and supervision
Utilities
Steam
Electricity
Heat credit
Maintenance'
Labor and material
Analyses
Total conversion costs
Total direct costs
Annual quantity
5,365 short tons
407 short tons
8,760 man-hr
5,978 MBtu
15,822,000 kWh
108,400 MBtu
2,920 man-hr
Unit cost, S
150.00/short ton
10,605.00/short ton
12.50/man-hr
2.00/MBtu
0.029/kWh
2.00/MBtu
17.00/man-hr
% of
Total annual annual revenue
cost, $ requirement
804,800
4,316,000
5,120,800
109,500
12,000
458,800
(216,800)
436,000
49,600
849,100
5,969,900
8.62
46.20
54.82
1.17 .
0.13
4.91
(2.32)
4.67
0.53
9.09
63.91
Indirect Costs
Capital charges
Depreciation, interim replacements, and
insurance at 6.03! of total depreciable
Investment
Average cost of capital and taxes at 8.6%
of total capital Investment
Overheads
Plant, SOX of conversion costs less utilities
Administrative, 10% of operating labor
Total indirect costs
Spent catalyst disposal
Total annual revenue requirements
1,234,400
1,822,300
297,600
11.000
3,365,300
6,000
9,341,200
13.2.1
19.51
3.19
0.12
36.03.
0.06
100.00
Equivalent unit revenue requirements for NO* removal
$/ton $/MBtu
Mills/kWh coal burned heat input
2.77
6.23
0.30
Basis
Midwest plant location, 1980 revenue requirements.
Remaining life of power plant, 30 yr.
Power unit on-stream time, 7,000 hr/yr.
Coal burned, 1,500,100 tons/yr, 9,000 Btu/kWh.
Investment and revenue requirement for disposal of fly ash excluded.
Total direct investment, $10,899,000; total depreciable investment, $20,574,000; and total capital invest-
ment, $21,190,000.
149
-------
TABLE 51. UOP SFGT-N, NOX-ONLY PROCESS
SUMMARY OF ANNUAL REVENUE REQUIREMENTS - REGULATED UTILITY ECONOMICS*
(500-MW new coal-fired power unit, 3.5% S in coal;
90% NOX removal)
% of
Total annual annual revenue
Annual auantltv Unit cost, $ cost, S requirement
Direct Costa
Raw materials
NH3
Catalyst
Total raw materials coats
Conversion costs
Operating labor and supervision
Utilities
Steam
Electricity
Heat credit
Maintenance
Labor and material
Analyses
Total conversion costs
Total direct costs
6,395 short tons 150.00/short ton 959,300
2,^49,000
3,408,300
8,760 man-hr 12.50/roan-hr 109,500
7,140 KEtu 2.00/MBtu 14,300
15,796,000 kWh 0.029/kWh 458,100
103,110 MBtu 2.00/MBtu (206,200)
393,800
2.920 man-hr 17.00/man-hr 49,600
619,100
4,227,400
13.35
34.07
47.42
1.52
0.20
6.33
(2.87)
5.48
0.69
11.40
58.82
Indirect Costa
Capital charges
Depreciation, interim replacements, and
insurance at 6.0% of total depreciable
investment
Average coat of capital and taxes at 8.6%
of total capital investment
Overheads
Plant, 50% of conversion costs less utilities
Administrative, 10% of operating labor
Total indirect costs
Spent catalyst disposal
Total annual revenue requirements
1,085,100
1.584,400
276,500
11,000
2,957,000
2,200
7.186,600
15.10
22.05
3.85
0.15
41.15
0.03
100.00
Equivalent unit revenue requirements for NOX removal
S/ton $/MBtu
Mills/kWh coal burned heat input
2.13 4.79 0.23
Basis
Midwest plant location, 1980 revenue requirements.
Remaining life of power plant, 30 yr.
Power unit on-stream time, 7,000 hr/yr.
Coal burned, 1,500,100 tons/yr, 9,000 Btu/kUh.
Investment and revenue requirement for disposal of fly ash excluded.
Total direct investment, $9,844,000; total depreciable investment, $18,085,000; and total capital investment,
$18,423,000.
150
-------
overheads, are 41.2% of the annual revenue requirement or about $3.0M. The
spent catalyst disposal cost is $2k, 0.03% of the annual revenue requirement.
NH3 is the raw material used in the largest quantity. The cost of NH3
is about 13.3% of the total annual revenue requirement, i.e., almost $0.96M
annually. The catalyst cost of $2.4M, the largest individual expense,
represents approximately 34.1% of the annual revenue requirements. Major
conversion costs are plant maintenance at $0.4M and electricity (primarily
for flue gas blowers) at $0.5M. A heat credit valued at about $0.2M was
applied based on the heat of reaction from the reduction of NOx with NH3
plus the heat released from the combination of excess NH3 with oxygen, less
the preheat required for the NH3 and air from the NH3 injection system,
The net heat released was considered to be recovered in the air heater.
The capital charges account for about 37.1% of the total annual revenue
requirements, i.e., about $2.7M. Overheads are 4.0% or $0.29M of the annual
revenue requirements. The spent catalyst disposal cost, based on disposal
cost for alumina, the catalyst carrier, is estimated to be $2k.
UOP SFGT-SN, SOx-NOx Process
The total annual revenue requirements for the UOP SFGT-SN process for
simultaneous SOx~NOx removal are approximately $22.5M as shown in Table 52.
This is equivalent to a unit revenue requirement of 6.64 mills/kWh. The
direct costs, comprising raw material and conversion costs are 67.0% of the
total annual revenue requirements or about $15M. Indirect costs comprising
capital charges and overheads, are 49.3% of the total annual revenue require-
ments. The spent catalyst disposal cost is $6k, only 0.03% of the annual
revenue requirement. A credit for byproduct sulfuric acid sales revenue of
$3.66M or 16.3% of the total annual revenue requirements is applied.
Naphtha is the raw material used in the largest quantity and represents
a raw material cost of $4.2M. Catalyst cost is the major raw material cost
of $6.6M. Major conversion costs are: 0) electricity costs (primarily
for flue gas blowers, sulfuric acid plant, and quench circulating pumps) of
about $2.0M; (2) plant maintenance cost of $1.5M; and (3) naphtha for fuel
costs of almost $1.1M. A heat credit valued at $2.2M is applied. The net
heat credit was based on the following:
1. Credit for heat reductions (oxidation of Cu, sulfation of CuO, and
reduction of CuS04 and CuO).
2. Credit for heat of reactions for NOX reduction with NH3 and combina-
tion of excess NH3 with oxygen.
3. Credit for combustion of purged gas.
4. Credit for vented gas combustion.
5. Credit for dewpoint suppression as a result of 803 removal.
6. Debit for NH3, air, and steam, superheat.
151
-------
TABLE 52. UOP SFGT-SN, SOX-NOX PROCESS
SUMMARY OF ANNUAL REVENUE REQUIREMENTS - REGULATED UTILITY ECONOMICS3
(500-MW new coal-fired power, unit, 3.5% S in coal;
90% NOX removal, 90% S02 removal)
Direct Costs
Raw materials
NH3
Naphtha
Catalyst
Steam-naphtha reformer catalyst
Total raw materials cost
Conversion costs
Operating labor and supervision
Utilities
Naphtha
Steam
Process water
Electricity
Heat credit
Maintenance
Labor and material
Analyses
Total conversion costs
Total direct costs
Annual quantity
6,395 short tons
8,400,000 gal
29,200 man-hr
2,100,000 gal
112,308 MBtu
2,134,000 kgal
70,550,000 kWh
1,121,680 MBtu
4,380 man-hr
Unit cost, $
150.00/short ton
0.50/gal
12. 50 /man-hr
0.50/gal
2.00/MBtu
0.06 /kgal
0.029/kUh
-2.00/MBtu
17. 00 /man-hr
% of
Total annual annual revenue
cost , $ requirement
959,300
4,200,000
6,648,000
100,000
11,907,300
365,000
1,050,000
224,600
128,000
2,046,000
(2,243,400)
1,492,900
74.500
3,137,600
15,044,900
4.27
18.69
29.59
0.45
53.00
1.62
4.67
1.00
0.57
9.11
(9.99)
6.65
0.33
13.96
• 66.96
Indirect Costs
Capital charges
Depreciation, Interim replacements, and
insurance at 6.0% of total depreciable
Investment
Average cost of capital and taxes*at 8.6%
of total capital investment
Overheads
Plant, 50% of conversion costs less utilities
Administrative, 10% of operating labor
Marketing, 10% of sales revenue
Total indirect costs
Spent catalyst disposal
Gross average annual revenue requirements
3,931,300
5,775,100
966,200
36,500
366.000
11,075,100
6,000
26.126,000
116.29
Byproduct Sales Revenue
Total byproduct sales revenue
Total annual revenue requirements
122,000 short tons -30.00/short ton
(3.660.000)
(3,660,000)
22.466.000
100.00
Equivalent unit revenue requirements for S02 and NOX removal
$/ton $/MBtu
Mills/kWb coal burned heat input
6.64 14,98 0.71
Basis
Midwest plant location, 1980 revenue requirements.
Remaining life of power plant, 30 yr.
Power unit on-stream time, 7,000 hr/yr.
Coal burned, 1,500,100 tons/yr, 9,000 Btu/kWh.
Investment and revenue requirement for disposal of fly ash ex:luded.
Total direct investment, $37,322,000; total depreciable Investment, $65,521,000; and total capital investment,
$67,153,000.
152
-------
The capital charges are the major expense and account for 43.2%, i.e.,
about $9.7M, of the total annual revenue requirement. Overheads are 6.1%
or about $1.4M, of the total annual revenue requirements. The spent catalyst
disposal cost, based on disposal cost for alumina, the catalyst carrier, is
estimated to be $6k.
The $3.66M credit for sulfuric acid sales revenue is based on the
annual sale of about 122,000 short tons of 100% H2S04-
Limestone FGD
The annual revenue requirement for the limestone FGD process with 90%
S02 removal, which is used in conjunction with each of the dry N0x-only
FGT processes for overall system comparability with SOX-NOX processes, was
estimated to be $14,651,300 (17).
ESP
The annual revenue requirements of the various ESP's were estimated by
the same procedures as given for the NOX FGT processes and are summarized
below,
1. The annual revenue requirement for four 99.5% efficient cold ESP's,
to be used with the Hitachi Zosen and UOP SFGT-N processes, was
estimated to be $2,220,000. The electricity consumption was based
on 2 watts/ft^ of collecting area.
2. The annual revenue requirement for four 99.5% efficient cold ESP's,
to be used with the UOP SFGT-SN process, was estimated to be
$2,983,500. The electricity consumption was based on 2 watts/ft^
of collecting area.
3. The annual revenue requirement for four 98% efficient hot ESP's, to
be used with the Kurabo Knorca process, was estimated to be
$2,687,300. The electricity usage was based on 4 watts/ft2 of
collecting area.
4. The annual revenue requirement for one 94% efficient cold ESP, to
be used with the Moretana Calcium process, was estimated to be
$1,497,200. The electricity usage was based on 2 watts/ft2 of
collecting area.
Overall Comparison
The estimated total capital investment and annual revenue requirements
for each of the NOx FGT alternatives based on the previously described
design and economic premises are shown in Tables 53 and 54 respectively.
The total capital investments range from $79*.6M ($165/kW) for the UOP SFGT-N
system to $203.6M ($<+82/kW) for the IHI SOx-NOx system.
153
-------
The three dry N0x-only processes, UOP, Kurabo, and Hitachi Zosen, have
total capital investments of $18.4M, $21.2M, and $23.3M respectively.
Including a total capital investment of $50.4M for the limestone FGD system,
and either $10.8M or $12.1M for an ESP, depending upon the NOx process, the
total capital investment for comparison of SOx-NOx-PM systems employing dry
NOx-only processes combined with limestone FGD and ESP is $79.6M, $83.7M,
and $84.5M for the UOP, Kurabo, and Hitachi Zosen systems respectively.
The UOP SFGT-SN process, the only dry S0x-N0x process, has a total
capital investment cost of about $67.2M. With the addition of $14.6M for PM
removal with ESP, the total capital investment for the SOX-NOX-PM system is
$81.8M (i.e., $169/kW). This investment is 2.4% greater than for the UOP
SFGT-N system (least costly SCR-FGD system), but 3.5% less than for the
Hitachi Zosen system (most costly SCR-FGD system).
TABLE 53. TOTAL CAPITAL INVESTMENT
FOR THE ALTERNATIVE SOx-NOx-PM SYSTEMS
Total capital
Process
UOP SFGT-N
UOP SFGT-SNb
Kurabo Knorca
Hitachi Zosen
Moretana Calcium
Asahi
IHI
FGT
18.4
67.2
21.2
23.3
88.0
104.9
203.6
FGD
50.4
—
50.4
50.4
-
-
—
investment, M$a
ESP
10.8
14.6
12.1
10.8
7.2
-
—
Total
79.6
81.8
83.7
84.5
95.2
104.9
203.6
$/kW
Total
165
169
174
175
205
233
482
$/aft3/min
Total
37.1
38.1
39.0
39.4
44.4
48.9
94.9
a. Basis
Midwest plant location represents project beginning mid-1977,
ending mid-1980. Average cost basis for scaling, mid-1979.
Stack gas reheat to 175°F by indirect steam reheat.
Minimum in-process storage; only pumps are spared.
Disposal pond located 1 mile from power plant.
Investment requirements for fly ash disposal excluded.
Construction labor shortages with accompanying overtime pay
incentive not considered.
b. Based on hydrogen production from naphtha and H2S04 byproduct from S02.
.The total capital investments of the wet SOX-NOX processes are
substantially higher and more divergent than those of the dry SCR-FGD
systems. With an ESP cost of $7.2M, the Moretana Calcium system at $95.2M
($205/kW) is approximately 24% higher than the least capital-intensive dry,
SCR-FGD system (UOP SFGT-N) while the Asahi system at $104.9M ($233/kW) is
an additional 14% higher than the Moretana Calcium system. (It should be
noted that to obtain 90% NOjj reduction, Moretana Calcium and Asahi processes
achieve 95% and 99% S02 removal which is greater than the 90% S02 removal
154
-------
of the other alternative FGT processes.) The IHI system has a total capital
investment of $203.6M ($482/kW), approximately double that of the Asahi
system and almost 3 times that of the least capital-intensive dry systems.
The total annual revenue requirements for SOX-NOX-PM control ranged
from 7.13 mills/kWh for the cheapest dry process (UOP SFGT-N) to more than
19.8 mills/kWh for the IHI process. When combined with total annual revenue
requirements for the limestone FGD system and an ESP, the revenue require-
ments for the three SCR-FGD systems ranged from 7.13 mills/kWh for the UOP
SFGT-N system to 8.60 mills/kWh for the Hitachi Zosen system. The revenue
requirement for the UOP SFGT-SN process, the only dry SOX-NOX system, was
6.6 mills/kWh (including a byproduct sales credit). With the addition of
the ESP, the revenue requirement for this SOX-NOX-PM system becomes 7.53
mills/kWh. This revenue requirement is only 6% higher than the least
expensive SCR-FGD system. Without the credit for sale of byproduct H2S04,
the revenue requirement for the UOP SFGT-SN process plus ESP becomes 8.60
mills/kWh. This is 21% higher than the least costly SCR-FGD system.
TABLE 54. TOTAL ANNUAL REVENUE REQUIREMENTS
FOR THE ALTERNATIVE SOX-NOX-PM SYSTEMS
Equivalent
Annual revenue revenue requirements,
requirements, M$a mills/kWh
Process
UOP SFGT-N
UOP SFGT-SNb
Kurabo Knorca
Hitachi Zoscn
Moretana Calcium
Asahi
IHI
FGT
7.2
22.5
9.3
12.2
38.1
39.8
58.6
FGD
14.7
14.7
14.7
—
-
ESP
2.2
3.0
2.7
2.2
1.5
-
Total
24.1
25.5
26.7
29.1
39.6
39.8
58.6
Total
7.13
7.53
7.91
8.60
12.20
12.63
19.82
Basis
Midwest plant location, 1980 revenue requirements.
Remaining life of power plant, 30 yr.
Power unit on-stream time, 7,000 hr/yr.
Coal burned, 1,500,100 tons/yr, 9,000 Btu/kWh.
Investment and revenue requirement for disposal of
fly ash excluded.
Based on hydrogen production from naphtha and ^50$ byproduct from S02
155
-------
The wet NOX FGT systems are significantly more expensive to operate
than the dry systems. The Asahi systems and Moretana Calcium systems at 12.6
mills/kWh and 12.2 mills/kWh, respectively, are the cheapest of the wet
processes but are still nearly twice as -expensive to operate as the less
expensive dry N0x-only - limestone FGD and simultaneous S0x-N0x systems.
In addition, the IHI system at 19.8 mills/kWh is significantly higher. Thus
it would appear from the total annual revenue requirements that the wet
SOX-NOX processes cannot economically compete with the dry SCR-FGD or
SOX-NOX processes under the premises used for this study. Under different
premises, however, it may be possible that these wet S0x-N0x processes might
be feasible, e.g., with retrofit if auxiliary heat were required with the
dry FGT alternatives, or if various byproduct credit situations were
considered.
Based on an accuracy range of -20% to +40%, the resulting range of
estimated capital investment and annual revenue requirement for the SOjj-NOx-
PM control systems is given in Table 55.
Energy Requirements for the Alternative NOY FGT Systems
Because of increasing concern about the energy requirements of various
processes, evaluation of energy costs is important. A comparison of the
energy requirements for the various alternative NOX FGT systems is shown in
Table 56. Fuel, steam, and heat credits are listed in equivalent MBtu/hr.
The electrical consumption is shown as equivalent MBtu/hr. The total
equivalent energy consumption is listed as a percentage of the boiler
capacity. The energy consumption for the limestone FGD system and the
various ESP's has been included.
The dry systems have the lowest energy requirement; energy equivalent
to 4% of the boiler capacity is consumed. The UOP SFGT-SN process requires
over 3% of the boiler capacity. The actual equivalent energy consumption,
however, has been reduced by about 160 MBtu/hr through the application of
the various heat credits claimed by UOP. Where applicable, similar heat
credits have been given to the other NOjj FGT processes. Combined with the
energy for an ESP, the energy for the UOP SFGT-SN system totals 3.6% of the
boiler capacity.
The dry N0x-only processes have low energy requirements, much less than
1% of boiler capacity in all three cases. Combined with the energy costs for
the limestone FGD system and the ESP, energy costs for the SCR-FGD systems
are still relatively low, ranging from 3.7 to 3.9% of boiler capacity.
The energy requirements for the wet systems range from 365 MBtu/hr
(8.1% of the boiler capacity) for the Moretana Calcium system, with an ESP,
to 836 MBtu/hr for the IHI system (18.6% of the boiler capacity). Thus the
wet simultaneous systems are 2 to nearly 5 times as energy intensive as
the dry NOx FGT alternatives. In addition, consuming more than 10% of the
new boiler capacity would severely affect new coal-fired boiler economics.
156
-------
TABLE 55. ACCURACY RANGE OF THE ESTIMATED
CAPITAL INVESTMENT AND ANNUAL REVENUE REQUIREMENT
FOR THE SOX-NOX-PM SYSTEMS
(Based on a -20% to +40% range)
Capital investment,
$/kW
Unit revenue
requirements,
mills/kWh
Process
UOP SFGT-N
UOP SFGT-SN3
Kurabo Knorca
Hitachi Zosen
Moretana Calcium
Asahi
IHI
Low
132
135
139
140
164
186
386
Base
165
169
174
175
205
233
482
High
231
237
244
245
287
326
675
Low
5.70
6.02
6.33
6.88
9.76
10.10
15.86
Base
7.13
7.53
7.91
8.60
12.20
12.63
19.82
High
9.98
10.54
11.07
12.04
17.08
17.68
27.75
a.
Based on hydrogen production from naphtha and
byproduct from S02«
157
-------
TABLE 56. COMPARISON OF ENERGY REQUIREMENTS FOR THE VARIOUS NOx FGT PROCESSES3
Ul
00
Heat Total equivalent
Process
Limestone FGD
Cold ESP (99.5% efficient)
Cold ESP (99.5% efficient -
UOP SFGT-SN)
Cold ESP (94% efficient)
Hot ESP (98% efficient)
Asahi
IHI
More tana Calcium
Moretana Calcium and ESP
(94%)
Hitachi Zosen
Hitachi Zosen and FGD and ESP
(99.5%)
Kurabo Knorca
Kurabo Knorca and FGD and ESP
(98%)
UOP SFGT-N
UOP SFGT-N and FGD and ESP
(99.5%)
UOP SFGT-SN
UOP SFGT-SN and ESP
(99.5%)
Fuel,
MBtu/hr
_
-
-
-
-
216
23
-
-
-
-
-
-
-
-
200
200
Steam,
MBtu/hr
78
-
-
-
-
105
232
117
117
1
79
1
79
1
79
18
18
Electricity,
MBtu/hr
74
8
13
4
19
189
581
244
248
20
102
20
113
20
102
91
104
credit, energy
con sump t ion, b
MBtu/hr % of boiler capacity
-
_^
—
-
-
-
-
-
14
14
16
16
15
15
160
160
3
0.2
0.3
0.1
0.4
11
19
8
8
0.2
4
0.1
4
0.2
4
3
4
a. Does not include energy requirement
b. Based on a 500-MW boiler,
and a boiler efficiency of
represented
a gross heat rate of
90% for
by raw materials.
9,000 Btu/kWh
for generation of
electricity,
generation of steam.
-------
Wet Process Gas-Phase Oxidation Versus Absorption of NOx
The IHI and Moretana S0x-N0x processes based on gas-phase oxidation were
devised to avoid absorption of relatively insoluble NO and thus reduce the
cost of the SOX-NOX absorption section. However, it appears that the choice
of oxidant dictates whether oxidation is more economical than absorption. The
ozone generating system cost is 38.3% of the total capital investment for the
IHI process. In fact, the direct capital investment for the ozone generating
system is nearly as much as the total capital investment of entire SCR-FGD
systems. The investment for the chlorine dioxide system in the Moretana
Calcium process is about -19% of the capital investment, a lesser but still
significant cost, reflecting the relatively large capital investment of the
IHI process. However, the Asahi .process, an absorption-reduction-type process,
has a higher capital investment than the Moretana Calcium process primarily
as a result of the capital costs for systems to absorb the NOX arid to
regenerate the sodium-based scrubbing solution.
The annual revenue requirements for both the Asahi and the Moretana
Calcium systems are 12.6 mills/kWh and 12.2 mills/kWh, respectively, while
IHI at 19.8 mills/kWh is significantly more expensive. The capital charges
are primarily responsible for the differences in the annual revenue require-
ments, $29.5M for IHI versus $12.5M for Moretana Calcium and $15.1M for
Asahi. Although the direct costs of each of these processes are essentially
the same at $22.OM to $25.5M for each, the annual raw material costs are
$1.8M for IHI, $6.4M for Asahi, and $12.9M for the Moretana Calcium process.
Annual conversion costs vary from $10.7M for the Moretana Calcium process
to $23.7M for IHI.
The use of an iron-chelating compound to assist the absorption of NO
rather than oxidizing NO to N0£ reduces capital investment because a less
expensive nitrogen byproduct section is required. The Asahi process which
uses Fe+2-EDTA converts most of the absorbed NOX to molecular nitrogen and
only relatively small amounts of complex nitrogen-sulfur compounds are formed.
The use of a gas-phase oxidant on the other hand results in the conver-
sion of the absorbed NOx to lesser amounts of molecular nitrogen and larger
amounts of either nitrate salts or a combination of nitrate salts and complex
nitrogen-sulfur compounds, depending on the gas-phase oxidant used. The use
of C102 as the gas-phase oxidant in the Moretana Calcium process results in
50% of the absorbed NOjj forming nitrate salts and the remaining 50% being
converted to molecular nitrogen. The IHI process using ozone results in only
10-15% of the absorbed NOX being converted to molecular nitrogen with most of
the remainder forming complex nitrogen-sulfur compounds.
The capital investment costs for the nitrogen treatment sections reflect
this inverse relationship between investment and the amount of NOX reduced
to molecular nitrogen in each process. The total capital investment of the
nitrogen treatment section of the Asahi process, which forms the higher
ratio of molecular nitrogen, is the lowest at slightly more than$1.17M,
while in the IHI process, which converts only 10-20% of the absorbed NOx to
molecular nitrogen, the total capital investment for the nitrogen treatment
section is about $5.15M.
159
-------
Thus it appears that wet processes based on oxidation of NOX with ozone
(IHI) are not economically competitive with the other wet FGT processes
primarily because of the costs involved in ozone generation. The use of C102
as the oxldant, as in the Moretana Calcium process, results in a more favor-
able economic comparison with the absorption-reduction Asahi process. The
Asahi process, which is about 14% higher in total capital investment than
the Moretana Calcium process and similar to it in revenue requirements, cannot
be definitely ruled out on the basis of this preliminary study. Under other
premises — a higher NOX concentration in the flue gas, for example — the
economic positions of these two processes could be reversed.
Wet SOx-NOx Versus Dry
From a comparison of total capital investments, the dry S0x-N0x system
(UOP SFGT-SN), at about $81. 8M ($169/kW) , is much less expensive than any
of the wet SOX-NOX systems, which range from 21% to 185% higher. The IHI
system, with 185% higher capital costs than the UOP system does not appear
economically competitive. Because of the preliminary nature of these
estimates, however, it is impossible to declare either the Asahi or the
Moretana Calcium system as economically unfeasible compared with UOP solely
on the basis of total capital investment.
A comparison of the annual revenue requirements for UOP and the wet
systems, however, indicates that the wet S0x-N0x FGT systems are not
economically competitive. Annual revenue requirements for the UOP system
are 7-53 mills/kWh with byproduct credit and 8.60 mills/kWh without this
credit whereas those for the Asahi, Moretana Calcium, and IHI systems, are
about 12.63, 12.20, and 19.82 mills/kWh respectively. With about 62% to
163% higher annual revenue requirements (42% to 130% without byproduct
credit for UOP system) and about 21% to 185% higher capital investment, it
would appear that the more expensive wet SOX-NOX systems would be eliminated
from further contention for most applications. However, it would be
difficult to extrapolate the result of these base case estimates to all
potential conditions.
Another factor to consider is the relatively early status of development
of the Asahi process. With the maturation of the Asahi technology it is
possible that the various costs could be reduced to make the process competi-
tive. However, it should also be pointed out that in most instances as the
technology has been further refined, the costs have tended to increase rather
than decrease,
Wet SOy-NOx Versus SCR-FGD
The total capital investments for the various combination SCR-FGD
systems, which ranged from $79. 6M <$165/kW) to $84. 5M ($175/kW) , were less
than the investment required for the various wet SOX-NOX systems. The wet
systems ranged from 24% to 192% higher than the UOP SFGT-N FGD system,
which had a total capital investment of $79. 6M. Based on comparison of
capital investments it appears that the Asahi and IHI systems are not
economically competitive when compared with the SCR-FGD systems under the
premises of this study.
160
-------
A comparison of the annual revenue requirements reinforces this view.
The annual revenue requirements for the three SCR-FGD systems range from
7.13 mills/kWh for UOP to 8.60 mills/kWh for Hitachi Zosen, while those for
the wet SOX-NOX systems range from 12.20 to 19.82 mills/kWh, 71 to 178%
higher than the UOP SFGT-N FGD system. Thus it would appear that for most
applications, the wet SOx-NOx systems (Asahi, IHI, and Moretana Calcium
processes) are not economically competitive with the SCR-FGD systems.
Dry SOy-NOy Versus SCR-FGD
The total capital investment of $81.8M (.$169/kW) for the UOP SOX-NOX
system is between the investment for SCR-FGD combination systems, which
range from $79.6M to $84.5M in total capital investment. Any difference is
well below the accuracy of these capital investment estimates so that there
is no significant difference for these two types of FGT within the premises
of this study.
The annual revenue requirements comparison has similar results. The
annual revenue requirements for the three SCR-FGD systems range from 7.13
mills/kWh for UOP to 8.60 mills/kWh for the Hitachi Zosen system while the
revenue requirements for the UOP SOX-NOX system are 7.53 mills/kWh with
byproduct credit and 8,60 mills/kWh without the credit. Thus, with the
byproduct credit the annual revenue requirement for the dry SOx-NQx removal
system is 6% higher than the UOP SFGT-N system. The other SCR-FGD systems
range from 5% (for Kurabo) to 14% (for Hitachi Zosen) higher than the UOP
SFGT-SN system. Without the byproduct credit, the UOP SFGT-SN system is 21%
and 9% higher than the UOP SFGT-N and Kurabo systems, respectively, and the
same as the Hitachi Zosen system. However, these results may vary under
different cases, for example, if the annual catalyst replacement on these
dry processes (which was a large contributing factor) were significantly
different. Also, for the cost estimate of the UOP SFGT-SN process, hydrogen
production from naphtha and conversion of S02 to H2S04 byproduct are
assumed. However, if the future supply of naphtha became uncertain, it may
be necessary to produce hydrogen by an alternative method, such as, coal
gasification. In addition, if there is no market for H2S04 byproduct, S02
processing alternatives may be used, for example, production of elemental
sulfur or liquid S02. Inclusion of alternative processing schemes for the
UOP SFGT-SN process may significantly alter these economic evaluations.
The sensitivity of the dry processes to the cost of NH3 has been the
subject of some concern. In this study the annual NH3 cost was less than
10% of the annual revenue requirement for two of the three dry NOx FGT
processes (Hitachi Zosen and Kurabo Knorca) and the NH3 cost accounted for
13% of the annual revenue requirement for the remaining dry process (UOP
SFGT-N). Thus, under the premises used in this study, the average annual
revenue requirement may be expected to increase about 10% when the NH3 cost
is doubled.
161
-------
Moving-Bed Reactor Versus Parallel Flow, Fixed-Bed Reactor
The capital investment and annual revenue requirements for the Kurabo
Knorca moving-bed reactor process are $43.9/kW and 2.77 mills/kWh, respec-
tively, for only the NOX removal portion. The capital investment and annual
revenue requirements for the parallel flow, fixed-bed reactor processes are
$38.1/kW and 2.13 mills/kWh for UOP SFGT-N, and $48.2/kW and 3.61 mills/kWh
for Hitachi Zosen. Thus, the costs for the moving-bed reactor fall between
the range of costs for the parallel flow, fixed-bed processes. Therefore, no
significant distinction can be made based on this level of study between
the moving-bed and parallel flow, fixed-bed processes. Actually, the major
economic differences among the N0x-only FGT processes are the catalyst costs.
The annual catalyst cost, based on a 1-year-guaranteed life and in 1980
dollars, ranged from about $2.5M for UOP to almost $6.9M for Hitachi Zosen.
162
-------
CONCLUSIONS AND RECOMMENDATIONS
CONCLUSIONS
Several conclusions may be derived from the data produced in this study.
The major conclusions are listed below. It should be remembered that these
conclusions apply to the results from our base case and premises involving a
new, 500-MW coal-fired unit. The results and conclusions may vary with dif-
ferent cases, e.g., smaller or larger plant, lower or higher sulfur content
in coal, different NOX and SOX removal efficiencies, or various credits for
byproduct sales. Also, these results are based on currently available
information and may have to be revised as future testing better defines the
FGT technology and its impact on design and operation of downstream equipment
Capital Investment
1. A dry SCR-FGD system (HOP SFGT-N) has the lowest capital investment.
The IHI SOX-NOX removal process has the highest investment.
2. The wet SOX-NOX processes based on ozone as a gas-phase oxidant are
prohibitively expensive for coal-fired units where NOx concentrations
are relatively high (e.g., 600 ppm) .
3.- The wet gas-phase oxidation process using C102 (Moretana Calcium
process) requires less investment than the other wet FGT processes.
The investments for the Asahi and IHI processes were 14 and 135%
higher, respectively, than the Moretana Calcium process.
4. The dry SOjc-NOx simultaneous removal system (UOP SFGT-SN) has a
lower investment than the wet, simultaneous S0x-N0x control systems.
The investment ranges from 21% higher for the Moretana Calcium system
to 185% higher for the IHI system, than the dry UOP SFGT-SN removal
system.
5. The dry N0x-only removal processes in combination with a limestone
FGD unit have lower investments than the wet, simultaneous SOX-NOX
removal systems. In comparison with the least expensive dry NOX-
only process and limestone FGD unit, the investment ranges from 24%
(for Moretana Calcium) to 192% (for IHI) higher for the wet S0x-N0x
removal systems.
6. The dry, simultaneous SOX-NOX control process has an essentially
equivalent investment as the combination of dry N0x-only removal pro-
cess. and limestone FGD unit. The UOP SFGT-SN process has a 2.4%
greater investment than the least expensive SCR-FGD system (UOP SFGT-N)
and 3.5% lower investment than the most expensive SCR-FGD system
(Hitachi Zosen) .
163
-------
7. At this preliminary level of study, there is no significant difference
in the investment among the dry N0x-only FGT processes. In combination
with the limestone PGD unit, the most expensive NOx-only control process
(Hitachi Zosen) has just a 6.0% higher investment than the least
expensive process (UOP SFGT-N) .
JAnnual Revenue Requirement
1. The UOP SFGT-N process has the lowest annual revenue requirement. The
IHI SOX-NOX removal process has the highest revenue requirement.
2. The revenue requirement is similar for two of the wet processes while
one process has substantially larger revenue requirements. The revenue
requirement is 3.5% and 62% higher for the Asahi and IHI processes,
respectively, than the Moretana Calcium process.
3. The revenue requirement is significantly less for the dry
removal system than the wet SOX-NOX control systems. The revenue
requirement ranges from 62% greater for the Moretana Calcium system
to 163% greater for the IHI system than for the dry UOP SFGT-SN system
(with byproduct credit). In comparison to the UOP SFGT-SN system without
byproduct credit, the wet systems have 42% to 130% higher revenue
requirements.
4. The SCR-FGD systems have significantly lower revenue requirements' than
the wet SOX-UOX control systems. The revenue tequirement ranges
from 71% to 178% higher for the Moretana Calcium and IHI systems,
respectively, than for the SCR-FGD system with the lowest revenue
requirement (i.e., UOP).
5. The revenue requirement for dry S0x-N0x removal is in the range of
revenue requirements for the SCR-FGD systems. With the byproduct
credit the UOP SFGT-SN removal system has a 6% higher revenue require-
ment than UOP SFGT-N process plus FGD unit. The other SCR-FGD systems
range from 5% (for Kurabo) to 14% (for Hitachi Zosen) higher in revenue
requirements than the UOP SFGT-SW system. Without the byproduct credit,
the UOP SFGT-SN system has 21% and 9% higher revenue requirement than
the UOP SFGT-N and Kurabo SCR-FGD systems, respectively, but has an
equal revenue requirement to the Hitachi Zosen system.
6. The Moretana Calcium process has the greatest annual raw material cost,
i.e., almost $12. 9M. The IHI process has the lowest annual raw
material cost of $1.8M.
7 . The least energy-intensive system is the UOP simultaneous SOX-NOX
control system with a net energy consumption of 162 MBtu/hr with the
heat credit (322 MBtu/hr without the heat credit) . The IHI system is
the most energy- intensive using 835.4 MBtu/hr.
8. For the dry NOx-only FGT processes, the revenue requirement would be
expected to increase about 10% as the NH3 cost is doubled.
164
-------
RECOMMENDATIONS
Additional preliminary- and definitive-level economic assessments of
control systems should be performed. The results and conclusions of this
NOX FGT study are based on a preliminary-level estimate for the base case
of treating flue gas from a new, 500-MW coal-fired power plant. It is very
possible that the results would be different with variations from the above-
mentioned base case and thus extrapolations of these results to other
conditions could lead to erroneous conclusions. Also, current and future
development testing will better define the FGT technology and its impacts
on downstream equipment, such as the air heater, FGD system, ESP or baghouse,
and flue gas fan. In addition, combustion modification for NOx emission
control, which has received the primary attention in the United States and
is considered the least expensive NOX reduction system, has not been
considered in this study. Therefore, additional economic evaluations,
including FGT and combustion modifications, should be conducted. The FGT
processes and combustion modification could be evaluated both independently
and in combination with each other. In addition, these control techniques
should be compared under several case variations to determine the sensitivity
of these technologies to various design and economic considerations.
Suggested items which should be varied include treatment capacity, sulfur
composition in coal, total NOX removal efficiency, raw material stoichiometric
ratios, and byproduct sales revenue.
165
-------
REFERENCES
1. Araki, T. Sumitomo Metal America, Inc., New York. Private communica-
tion, March 1978.
2. Barrier, J. W., H. L. Faucett, and L. J. Henson. Economics of Disposal
of Lime/limestone Scrubbing Waste; Sludge-Fly Ash Blending and Gypsum
Systems. (TVA Bull. Y-140; EPA-600/7-79-069), February 1979.
3. Bauerle, G. L., S. C. Wu, and K. Nobe. Catalytic Reduction of Nitric
Oxide with Ammonia on Vanadium "Oxide arid Iron-Chromium Oxide. Ind. Eng.
Chem., Prod. Res. Develop., 14(4):268-273, 1975.
4. Chakrabarti, G., and C. Chu. Reduction of Nitric Oxide with Ammonia on
Copper Chromite and Nickel - Copper Oxide Catalyst. Atmos. Environ.,
6:297-307, 1972.
5. Cotter, J. E. TRW, Inc., Redondo Beach, California. Private communica-
tion, May 1978.
6. Faucett, H. L., J. D. Maxwell, and T. A. Burnett. Technical Assessment
of NOx Removal Processes for Utility Application. TVA Bull. Y-120;
EPA-600/7-77-127 (NTIS PB 276 637/6WP), November 1977.
7. Federal Power Commission. "Hydroelectric Power Evaluation," FPC P-35
(1968) and Supplement No. 1, FPC P-38 (1969). U.S. Government Printing
Office, Washington, DC.
8. Ikawa, K. Hitachi Shipbuilding and Engineering Company, Ltd., New York.
Private communication, February 1978.
9. Koutsoukas, E. P., J, L. Blumenthal, M. Ghassemi, and G. Bauerle.
Assessment of Catalyst for Control of HOx frfem^Stationary Power Plants,
Phase I, Vol. I. U.S. Environmental Protection Agency, EPA-650/2-75-
OOlb (NTIS PB 239 745), January 1975,
10. Markvart, M., and V. Pour. Selective Reduction of Nitrogen Oxides with
Ammonia. Chem. Prum., 19(1):8-12, 1969.
166
-------
11. Matsumoto, H. Ishikawajima-Harima Heavy Industries Company, Ltd.,
Tokyo. Private communication, February 1978.
12. McCutchen, G. D. "NOX Emissions Trends and Federal Regulation." Chem.
Eng. Prog., 73(8):58-63, August 1977.
13. Ohlmanziek, D. Emery Industries, Cincinnati, Ohio. Private communica-
tion, March 1978.
14. Peters, M. S., and K. D. Timmerhaus. Plant Design and Economics for
Chemical Engineers, 2d Ed., McGraw-Hill, New York, 1968,
15. Ricci, L. J. "EPA Sets Its Sights on Nixing CPI's NOX Emissions."
Chem. Eng. 84(4):32-36, February 14, 1977,
16. Terasaki, I. Asahi Chemical Industry Company, Ltd., Tokyo. Private
communication, March 1978.
17. Tomlinson, S. V., R. L. Torstrick, F. A. Sudhoff, and F. M. Kennedy.
Definitive SOx Control Process Evaluations;. Limestone, Double Alkali,
and Citrate FGD Processes. TVA-ECDfc-B-4; EPA-600/7-79-177, August 1979.
18. U.S. Environmental Protection Agency. "Standards of Performance for
New Stationary Sources." Fed. Regist., 36(247):24, 879, December 23,1971,
19. U.S. Environmental Protection Agency. Monitoring and Air Quality Trends
Report, 1974. EPA-450/1-76-001 (NTIS PB 252 269), February 1976.
20. U.S. Environmental Protection Agency. "New Stationary Source Performance
Standards: Electric Utility Steam Generating Units." Fed. Regist.,
44(113):33580-33624, June 11, 1979.
21. Verson, R. L. Universal Oil Products, Inc., Des Plaines, Illinois.
Private communication, March 1978.
167
-------
TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
1. REPORT NO.
EPA-600/7-80-021
2.
3. RECIPIENT'S ACCESSION NO.
4. TITLE AND SUBTITLE
Preliminary Economic Analysis of NOx Flue Gas
Treatment Processes
5, REPORT DATE
February I960
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S>
J.D.Maxwell, T.A.Burnett, and H.L.Faucett
B. PERFORMING ORGANIZATION REPORT NO.
TVA ECDP B-6
EPRI FP-1253
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Tennessee Valley Authority
Office of Power
Muscle Shoals, Alabama 35660
10. PROGRAM ELEMENT NO.
INE624
11. CONTRACT/GRANT NO.
EPA IAG-D8-E721-FU
12. SPONSORING AGENCY NAME AND ADDRESS
EPA, Office of Research and Development
Industrial Environmental Research Laboratory
Research Triangle Park, NC 27711
13. TYPE OF REPORT AND PERIOD COVERED
Final: 7/77 - 10/79
14. SPONSORING AGENCY CODE
EPA/600/13
15. SUPPLEMENTARY NOTES lERL-RTP project officer is J.D. Mobley, MD-61, 919/541-2915.
EPRI project monitor is Navin Shah.
IB. ABSTRACT
repOrt gives results of z. preliminary economic evaluation of seven
flue gas treatment (FGT) processes for the removal of NOx from power, plant flue
gas. The base case was a new, 500-MW power plant burning 3. 5% sulfur coal and
emitting 600 ppm NOx in the flue gas. Total capital investments and annual revenue
requirements for three dry NOx-only removal processes were $38-48/kW and 2.1-
3. 6 millsAWh, respectively. A dry process to simultaneously remove both SOx and
NOx had a capital cost of ^139/kW and an annual revenue requirement of 6. 6 mills/
kWh. Wet SOx/NOx systems, which also remove particulates , had capital costs and
annual revenue requirements of #205-482/kW and'12.2-19. 8 mills/kWh, respectively.
To permit comparisons between the alternative NOx control systems , the costs of
both a limestone slurry flue gas desulfurization unit and an ESP were added to the
cost of the dry NOx-only processes , and the cost of an ESP was added to the cost of
the dry SOx/NOx processes. Total capital investments for. these combined systems
based on dry NOx removal (both NOx-only and SOx/NOx removal) were about #165-
175/kW. Unit revenue requirements for the combined systems were 7. i-8.6 mills/
kWh for the dry processes. For the base case, wet FGT processes do not appear to
be economically competitive with either the dry NOx-only or SOx/NOx systems.
7.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lOENTIFIERS/OPEN ENDED TERMS
c. COSATI Field/Group
Pollution
Economic Analysis
Flue Gases
Nitrogen Oxides
Sulfur Oxides
Coal
Combustion
Desulfurization
Pollution Control
Stationary Sources
Flue Gas Treatment
13B
05C
21B
07B
21D
07A,07D
18. DISTRIBUTION STATEMENT
Release to Public
19. SECURITY CLASS (This Report)
Unclassified
21. NO. OF PAGES
198
20. SECURITY CLASS (Thispage)
Unclassified
22. PRICE
EPA Form 2220-1 (9-73}
168
------- |