c/EPA
United States
Environmental Protection
Agency
Air Pollution Training,Institute «
MD20
Environmental Research Center
- Research Triangle Park, NC 27711
EPA 450/2-81-005
December, 1981
Air
APTI
Course 415
Control of Gaseous
Emissions
Student Manual
-------
United States
Environmental Protection
Agency
Air Pollution Training Institute
MD20
Environmental Research Center
Research Triangle Park, NC 27711
EPA 450/2-81-005
December, 1981
Air
APTl
Course 415
Control of Gaseous
Emissions
Student Manual
Written by:
Gerald T. Joseph, P.E.
David S. Beachler
Northrop Services, Inc.
P.O. Box 12313
Research Triangle Park, NC 27709
Under Contract No.
68-02-2374
EPA Project Officer
R. E. Townsend
United States Environmental Protection Agency
Office of Air, Noise, and Radiation
Office of Air Quality Planning and Standards
Research Triangle Park, NC 27711
U 0. D-y:rc""-"yi!:r! frct-clion Agency
; J,, .:.-.--, • • :TI Street
Chicci/o, iiiinois 6JG04 M£f/-"'
-------
Notice
This is not an official policy and standards document. The opinions and selections
are those of the authors and not necessarily those of the Environmental Protection
Agency. Every attempt has been made to represent the present state of the art as
well as subject areas still under evaluation. Any mention of products or organiza-
tions does not constitute endorsement by the United States Environmental Protec-
tion Agency.
Availability
This document is issued by the Manpower and Technical Information Branch,
Control Programs Development Division, Office of Air Quality Planning and Stan-
dards, USEPA. It was developed for use in training courses presented by the EPA
Air Pollution Training Institute and others receiving contractual or grant support
from the Institute. Other organizations are welcome to use the document.
This publication is available, free of charge, to schools or governmental air
pollution control agencies intending to conduct a training course on the subject
covered. Submit a written request to the Air Pollution Training Institute, USEPA,
MD 20, Research Triangle Park, NC 27711.
Others may obtain copies, for a fee, from the National Technical Information
Service (NTIS), 5825 Port Royal Road, Springfield, VA 22161.
Sets of slides and films designed for use in the training course of which this
publication is a part may be borrowed from the Air Pollution Training Institute
upon written request. The slides may be freely copied. Some films may be copied;
others must be purchased from the commercial distributor.
U,3. environmental Protection Ageffey
n
-------
Table of Contents
Page
Chapter 1. Air Pollution Control 1-1
Overview 1-1
Gaseous Emission Control Techniques 1-3
Chapter 2. Basic Concepts of Gases 2-1
Expression of Gas Temperature 2-1
Expression of Gas Pressure 2-2
Density 2-4
Molal Units 2-5
Viscosity 2-5
Specific Heat 2-7
Relationships of Ideal Gases 2-7
Dalton's Law of Partial Pressure 2-9
Gas-Liquid Relationships 2-10
Reynolds Number 2-10
Chapter 3. Combustion 3-1
Introduction 3-1
Combustion Principles 3-1
Combustion Calculations 3-8
Combustion Equipment Used for Control of Gaseous Emissions 3-15
Introduction 3-15
Direct Combustion or Flaring 3-16
Incinerators 3-19
Catalytic Oxidation 3-27
Process Boilers Used as Incinerators 3-30
Heat Recovery Systems 3-31
Chapter 4. Absorption 4-1
Introduction 4-1
Mechanism of Absorption 4-2
Solubility 4-3
Absorption Design Theory 4-8
Design Procedures 4-13
Material Balance 4-14
Determining the Liquid Requirement 4-17
Sizing the Absorber 4-22
Absorption Equipment 4-37
111
-------
Page
Chapter 5. Adsorption 5-1
Introduction 5-1
Theory of Adsorption 5-2
Adsorption Equilibrium Relationships 5-7
Adsorbent Materials 5-12
Dynamic Adsorption Process 5-16
Factors Affecting Adsorption 5-18
Adsorbent Regeneration Methods 5-22
Adsorption Control Systems 5-26
Chapter 6. Condensation 6-1
Introduction 6-1
Condensation Principles 6-1
Condensers 6-3
Design of Condensers 6-9
Chapter 7. Control of Nitrogen Oxide Emissions 7-1
Introduction 7-1
Formation of Nitrogen Oxides in Combustion Sources 7-4
Combustion Modifications 7-5
Flue Gas Treatment 7-13
Other NOX Reduction Techniques 7-24
Chapter 8. Control of Sulfur Oxide Emissions 8-1
Introduction 8-1
Flue Gas Desulfurization 8-6
Nonregenerable FGD Processes 8-7
Lime Scrubbing 8-7
Limestone Scrubbing 8-11
Double Alkali Scrubbing 8-15
Regenerable FGD Processes 8-17
Wellman-Lord 8-18
Magnesium Oxide Process 8-21
Citrate Process 8-22
Dry SO2 Control Systems 8-23
Spray Dryer with a Baghouse or ESP 8-24
Dry Injection 8-27
Other Dry SO2 Processes 8-28
Comparison of FGD Systems 8-28
IV
------- An error occurred while trying to OCR this image.
-------
Figures
Figure p^
2-1 Comparison of Fahrenheit, Celsius, and absolute temperature scales 2-2
^-2 Gas-pressure relationship
2-3 Shearing stress in a moving fluid 95
3-1 Available heat for some typical fuels (referred to 60°F). . 3.10
3-2 Heat terms
3-3 Heat balance ' ' ' ' I
3-4 Smokeless flare tip 317
3-5 Ground flare
3-6 Typical thermal incinerator (UOP raw gas burner)...... . 3.19
3-7 Effects of temperature and residence time on rate
of pollutant oxidation.
3-20
3-21
3-8 Distributed (line) burner
> and hot combustion products
.3-24
3-23
waste gas and hot combiisrirm nrr,Hi,^o
3-10 Multijet burner
3-9 Mixing plate for waste gas and hot combustion products . . . . . . . . . . . . . 3.23
3-11 North American flame grid burner » 9,
3-12 Discrete burner *'*
3-13 Bridge wall baffle „
3-14 Ring and disc baffle
3-15 Typical honeycomb catalysts (metallic or ceramic) * 9«
3~i c /~i I . . . /
-------
Figure °
4-18 Cyclonic spray scrubber 4'4"
4-19 Typical venturi configuration 4~4;i
4-20 Spray venturi
4-21 Adjustable throat venturi 4"45
• 4 46
4-22 Jet or ejector venturi ^ "
4-23 Countercurrent packed tower 4"47
4-24 Cross flow packed tower • 4"48
4-25 Three bed cross flow packed tower 4'49 .
4-26 Common packing materials • 4"50
4-27 Liquid feed distributors 4"52
4-28 Plate tower 4"53
4-29 Scrubbing plates or trays 4"54
4-30 Flooded bed absorber 4"56
4-31 Fluidized bed absorber 4'57
4 ^R
4-32 Mist eliminators '•
5-1 Vapor adsorbed into pores of activated carbon 5-1
5-2 Mechanism of adsorption 5-2
5-3 Physical forces causing adsorption 5-5
5-4 Adsorption isotherm for carbon tetrachloride on activated carbon 5-8
5-5 Adsorption isoteres of H2S on 13X molecular sieve loading
in% weight 5"11
5-6 Adsorption isobar for benzene on carbon (benzene at 10.0 mm Hg) .... 5-12
5-7 Molecular screening in pores of activated carbon 5-15
5-8 Breakthrough curve 5"16
5-9 Carbon capacity vs. temperature 5-19
5-10 Pressure drop vs. flow rate through granular carbon beds 5-21
5-11 Two bed adsorption system 5-25
5-12 Thin bed adsorber: nine cell system 5-27
5-13 Pleated thin bed 5'28
5-14 Canister 5"28
5-15 Canister 5"29
5-16 Three bed system 5'30
5-17 Three bed vertical system 5-31
5-18 Horizontal bed 5"32
5-19 Three bed horizontal system 5-33
5-20 Rotary bed system 5-34
5-21 Fluidized bed adsorber (schematic) 5-35
6-1 Typical vapor pressure curve "'2
6-2 Direct contact condensers: (a) spray, (b) jet ejector,
and (c) barometric °"4
fi fi
6-3 Single-pass condenser °"°
6-4 Simplified air flow in multipass exchangers 6-7
6-5 Extended surface tubes 6'8
vn
-------
Figure page
6-6 Resistance to heat transfer around a cooling tube ................... 6-11
6-7 Temperature profiles in a heat exchanger for countercurrent flow,
cocurrent flow, and isothermal condensation with
countercurrent flow .
7- 1 Manmade nitrogen oxide emissions ................................ 7-2
7-2 Typical boiler ......................................... 7_g
7-3 Excess air ......................................... 7.7
7-4 Staged combustion .................................... 7_g
7-5 Over fire air ................................................. 7.9
7-6 Burners out of service .................................. 7. 10
7-7 Flue gas recirculation ....................................... 7_U
7-8 Schematic of a low NO* burner .................................. 7-12
7-9 Exxon Thermal DeNO* process .................................. 7-14
7-10 Selective Catalytic Reduction (SCR) process ....................... 7-16
7-11 Fixed bed reactor ...................................... 7. 17
7-12 Mobile bed reactor ........................................ 7_lg
7-13 Parallel flow reactor ................................... 7_lg
7-14 Parallel flow catalysts .......................................... 7. 19
7-15 Unit cell detail of catalyst grid .................................. 7-19
7-16 Shell UOP process ................................ '.'..'.'.'.'.'.'.'.'.'.'.7-21
7-17 Catalyst regeneration (Shell UOP) ............................... 7-22
7-18 Detail of catalyst (Shell UOP) ................................. ] .1-2$
8-1 Sulfur oxide emission sources ..................................... g.l
8-2 Scrubber and absorber train for lime scrubbing system ................ 8-8
8-3 Electrostatic precipitator and mobile bed scrubber train
for lime scrubbing system ...................................... g.g
8-4 Limestone scrubbing system .................................... g_12
8-5 Dual alkali scrubbing system .................................... g-15
8-6 Wellman-Lord system ....................................... 8-19
8-7 Spray dryers with baghouse ..................................... g_25
8-8 Spray dryer .................................................. g_25
9-1 Typical hoods and their entry coefficients .......................... 9-2
9-2 Flow diagram of air being pulled into a plain opening ................ 9-3
9-3 Effect of flange on air flow into a plain opening ...................... 9-4
9-4 Dimensions used in designing high canopy hoods controlling
emissions from hot sources ..................................... g.g
9-5 Open top tank degreaser with lip exhaust .......................... 9-12
9-6 Minimum ventilation rates for lip exhaust hoods .................... 9-13
9-7 Pressure terms used to describe air flow in a duct
(measured in inches) ......................................... 9-14
9-8 Friction loss chart ............................................. 9-17
9-9 Layout for Example 9-4 ........................................ 9-19
via
-------
Page
Fieure s
• 9 24
9-10 Typical point of operation "
9-11 Axial fans 9'25
9-12 Centrifugal fans 9"26
9-13 Typical performance curve for a backward curved
centrifugal fan 9"^y
9-14 Use of performance curves for fan selection 9'30
D-l Excess air concentrations of CO2 and O2 in stack D-3
E-l Prices for thermal incinerators without heat exchangers E-2
E-2 Prices for thermal incinerators with primary heat exchanger E-2
E-3 Catalytic incinerator prices E'2
E-4 Spray chamber costs vs. inlet gas volume E'3
E-5 Fabricated cost of carbon steel vessel E'4
E-6 Skirt and support costs for carbon steel vessel E-5
E-7 Cost of nozzles E"5
E-8 Cost of tray, support plate, or distributor E'6
E-9. Prices for packaged stationary bed carbon adsorption units
with steam regeneration E'8
E-10 Prices for custom carbon adsorption units E'9
E-11 Installed capital costs vs. condenser area for a complete
Kin
condenser section : ^"1U
1-1 Psychrometric chart. Properties of air and water-vapor
mixtures from 32° to 600°F I'2
Table
1-1
1-2
3-1
3-2
3-3
4-1
4-2
4-3
4-4
Tables
National primary ambient air quality standards
for gaseous pollutants
Typical gaseous pollutants and their sources ....
Combustion constants and approximate limits of flammability
of gases and vapors in air
Heat contents of various gases
Typical waste gas incinerators' operating temperatures (°F)
Partial pressure of SO2 in aqueous solution, mm Hg
Henry's Law constants for gases in H2O
Packing data
Empirical constants for Equation 4-26
Page
1-1
.1-4
.3-5
.3-9
3-20
.4-4
, .4-6
.4-24
.4-32
IX
-------
Table D
Page
5-1 Summary of characteristics of chemisorption and physical adsorption ... 5-7
5-2 Physical properties of major types of adsorbents 5.14
5-3 Regeneration of one pound of activated carbon loaded
with 20% ether 5 23
5-4 Adsorption filters r ao
6-1 Typical overall heat transfer coefficients in tubular
heat exchangers fi 19
7-1 New Source Performance Standards for NO* emissions from
fossil-fuel-fired steam generators rated at greater than
250 x 106 Btu/hr or 73 MW (thermal) heat input 7.3
8-1 New Source Performance Standards for SO2 emissions from
fossil-fuel-fired steam generators rated greater than
250 x 106 Btu/hr or 73 MW (thermal) heat input 8-3
8-2 Selected lime scrubbing FGD systems 8-10
8-3 Selected limestone scrubbing FGD systems 8-13
8-4 Selected dual alkali FGD systems 8-17
8-5 Wellman-Lord installations in the United States 8-20
8-6 Magnesium oxide installations in the United States 8-22
8-7 Key features of dry flue gas desulfurization systems 8-24
27
8-8 Commercial spray dryer FGD systems 8-
8-9 Cost estimates of selected FGD systems .......................... 8-29
9- la Range of capture velocities and duct velocities ....................... 9.5
9-lb Recommended minimum duct velocities .................. 9.5
Control volumes for low canopy hoods .................... 9. JQ
9-2
9-3 Equivalent resistance in feet of straight pipe ....................... 9-18
9-4a " "
9-4b Exhaust system calculations ............................. 9.23
9-5 Typical fan rating table .................................. 9.3!
Exhaust system calculations (trial figures) 9.21
E-1 Minimum shell thickness at ambient temperature (carbon steel) E-6
E-2 Cost of tower packing £.7
E-3 Additional costs for fabricator's engineering, purchasing,
administration and profit £.7
E-4 Technical assumptions for estimation of direct operating costs E-9
-------
Nomenclature
Symbols, metric units (English units)
a surface area available for absorption per unit volume of tower,
nWm3 (ftVft3)
A area, m2 (ft2)
Ac area of rising column of a gaseous emission, m2 (ft2)
A, top surface area of a hot emission source, m2 (ft2)
AHP air horsepower, kW (hp)
BHP brake horsepower, kW (hp)
c concentration, g/cm3 (lb/ft3)
CAI concentration of component A at the gas-liquid interface, g/cm3 (lb/ft3)
CAL concentration of component A in the liquid phase, g/cm3 (lb/ft3)
C empirical constant in BET equation, dimensionless
CB breakthrough capacity, weight %
C. coefficient of air entering an orifice, dimensionless
Cp specific heat, J/kg« °C (Btu/lb- °F)
Cp average specific heat, J/kg» °C (Btu/lb • °F)
C, saturation capacity of adsorber, weight %
d diameter of a pipe or duct, m (ft)
D adsorber bed depth, m (ft)
d, diameter of a tower or column, m (ft)
dw diameter of a fan wheel, m (ft)
Dc diameter of a column of hot rising fume, m (ft)
Dh diameter of a hood, m (ft)
D, diameter of a hot emissions source, m (ft)
E heat transfer effectiveness, %
£ mechanical efficiency of a fan, %
f percent of flooding
F absorber packing factor, m2/m3 (ftVft3)
AF free energy change, J/mol (Btu/mol)
gc gravitational constant, kg»m/N«s2 (Ib»ft2/lb/»sec)
G gas flow rate, kg/s (Ib/min)
G' superficial gas flow rate, kg/s»m2 (lb/min»ft2)
Gm gas molar flow rate, kg mol/s (Ib mol/sec)
Henry's Law constant, kPa/mol fraction (atm/mol fraction)
H enthalpy, J/g (Btu/lb)
Hc heat of combustion, J/m3 of fuel (Btu/ft3 of fuel)
H,4 available heat, J/m3 (Btu/ft3)
HOG height of a transfer unit based on overall gas phase, m (ft)
HOI height of a transfer unit based on overall liquid phase, m (ft)
hp perimeter of canopy hood, m (ft)
H, sensible heat, J/g (Btu/lb)
XI
-------
Hv latent heat of vaporization, J/g (Btu/lb)
HVC gross heating value, J/m3 (Btu/ft3)
HVN net heating value, J/m3 (Btu/ft3)
kg individual mass transfer coefficient based on gas phase,
kg mol/h»m2»kPa (Ib mol/hr»ftz»atm)
kf individual mass transfer coefficient based on liquid phase,
kg mol/h«m2»kPa (Ib mol/hr»ft2«atm)
k, individual mass transfer coefficient for combined surface migration and
pore diffusion, kg mol/h«m2»kPa (Ib mol/hr»ftz»atm)
KOG overall mass transfer coefficient based on gas phase,
kg mol/h»m2»kPa (Ib mol/hr.ft2«atm)
K0£ overall mass transfer coefficient based on liquid phase,
kg mol/h»m2»kPa (Ib mol/hr»ft2«atm)
/ length, m (ft)
L liquid flow rate, kg/s (Ib/sec)
L' superficial liquid flow rate, kg/s»m2 (lb/sec»ft2)
Lm liquid molar flow rate, kg mol/s (Ib/sec)
m slope of a line
m mass flow rate, kg/h (Ibs/hr)
M mass, kg (Ib)
Mat mass of theoretical air per unit mass of fuel combusted, kg/kg (Ib/lb)
MW molecular weight, g mol (Ib mol)
n number of moles
NX mass flux, kg mol/s»m2 (Ib mol/sec«ft2)
NOG number of transfer units based on overall gas coefficient
N0i number of transfer units based on overall liquid coefficient
Np number of theoretical plates
p partial pressure, kPa (psi)
p* partial pressure at equilibrium, kPa (psi)
p° vapor pressure, kPa (psi)
P* gage pressure, kPa (psig)
ps static pressure, kPa (psi)
PV velocity pressure, kPa (psi)
PAI partial pressure of component A at the gas-liquid interface, kPa (psi)
PAG partial pressure of component A in the gas phase, kPa (psi)
Pr total pressure, kPa (psi)
Pb barometric or atmospheric pressure, kPa (psi)
P absolute pressure, kPa (psi)
,JJ> power, kW (hp)
Q, volumetric flow rate, mVmin (cfm)
q heat rate, J/h (Btu/hr)
r radius, m (ft)
rpm fan speed (revolutions per min)
xn
-------
R ideal gas constant (see page 2-8)
Re Reynolds Number, dimensionless
T temperature, °C (°F)
AT,m log mean temperature difference, °C (°F)
ATm mean temperature difference, °C (°F)
U heat transfer coefficient, kj/h/m2 (Btu/hr/ft2)
v velocity, m/s (ft/sec)
V volume, m3 (ft3)
Va( volume of theoretical air required to combust a unit volume of fuel,
mVm3 (ftVft3)
w width, m (ft)
x mole fraction of contaminant in liquid phase
x distance, m (ft)
X mole fraction of contaminant in pure liquid
Xc sum of the distance from the hot source to the hood face plus
the distance from the top of the heat source to a hypothetical
vortex of the fume column, m (ft)
Xs degree of saturation in the mass transfer zone, percent
y mole fraction of contaminant in gas phase
Y mole fraction of contaminant in carrier gas
Yc distance from hood face to hot emissions source, m (ft)
Z height of packing in a tower, m (ft)
Zc distance from top of heat source to hypothetical vortex of fume column,
m (ft)
Greek letters
e density, kg/m3 (lb/ft3)
T shearing stress, kPa/m2 (lb/ft2)
v kinematic viscosity, mVs (ftVsec)
/j, viscosity, centipoise, Pa»s (centipoise)
0 residence time, s (sec)
-------
Chapter 1
Air Pollution Control
Overview
Controlling the emission of pollutants from industrial sources is important in pro-
tecting the quality of the air. Air pollutants can exist in the form of paniculate
matter or as gases. This text will focus on controlling gaseous air pollutants.
National Ambient Air Quality Standards (NAAQS) currently exist for five gaseous
pollutants. Table 1-1 lists the primary standards; note that these are ambient stan-
dards and not source emission limits.
Table 1-1. National primary ambient air quality standards for gaseous pollutants.
Pollutant
Sulfur dioxide, SO 2
Carbon monoxide, CO
Ozone, O$
Nonmethane
hydrocarbons, NMHC
Nitrogen dioxide, NO2
Standards
Annual: arithmetic mean
80 /ig/m3
(0.03 ppm)
—
—
100 /tg/m3
(0.05 ppm)
Other
365 /ig/ms (0.14 ppm)
24 hr max. once a yr.
40,000 /ig/m3 (35 ppm)
1 hr max. once a yr.
10,000/ig/m3 (9 ppm)
8 hr max. once a yr.
235 /ig/m3 (0.12 ppm)
1 hr max. once a yr.
160 /tg/m3 (0.24 ppm)
3 hr max. (6 to 9 a.m.)
once a yr.
—
Air cleaning devices have been reducing gaseous emissions from various
industrial sources for many years. Originally, air cleaning equipment was used only
if the gaseous contaminant was highly toxic or had some recovery value. Now with
legislation such as the 1970 Clean Air Act and the 1977 Clean Air Act Amend-
ments, control technologies have been upgraded and more industrial sources are
regulated in order to meet the NAAQS. In addition, State and local air pollution
agencies have adopted regulations that are in some cases more stringent than
Federal emission standards.
1-1
-------
The Clean Air Act as amended in 1977 contains basically three classes of control
technology which are aimed at reducing emissions from existing sources new
sources, and modified sources. These levels of control are: Reasonably Available
Control Technology (RACT), Best Available Control Technology (BACT), and
controls that reflect the Lowest Achievable Emission Rate (LAER).
In order to understand the distinctions between these various types of control
technologies it is necessary to become familiar with their regulatory definitions.
• First consider the following definitions from the U.S. Code of Federal
Regulations:
A new source is one which is contracted and installed at a facility after the date
emission standards are proposed for that industry.
Modification means any physical change or change in the operation of an
existing facility which increases the amount of an air pollutant emitted, or results
in the emission of an air pollutant not previously emitted into the atmosphere to
which a standard applies.
The term existing source refers to an air pollution source which was constructed
before the proposal date of the emission standard.
• On this basis, then:
Reasonably Available Control Technology is control technology applied to
existing sources to meet ambient air quality standards. This technology is generally
designated in the State Implementation Plan (SIP) so as to reduce emissions to
achieve ambient standards. RACT considers both the cost and technology available
for emission control. This control is usually less stringent than both BACT or
LAER.
Best Available Control Technology is control technology applied to new and
modified sources. Prevention of Significant Deterioration (PSD) is the review
process for sources in attainment areas and is required for each pollutant subject to
regulation under the Act. An attainment area is an air quality control region that
meets the National Ambient Air Quality Standards.
New Source Performance Standards (NSPS) are set for a number of source
categories. Best Available Control Technology must reduce emissions at least as
much as the NSPS limit. New Source Performance Standards that have been
promulgated are published first in the Federal Register and then in the U.S. Code
of Federal Regulations. BACT applied to a source must be technologically feasible
and must also reflect considerations of cost and energy usage. The BACT require-
ment is to be applied on a case-by-case basis by EPA regional engineers.
Lowest Achievable Emission Rate refers to control technology applied to new and
modified sources in a nonattainment area. A nonattainment area is an area or -
region where the National Ambient Air Quality Standard for a particular pollutant
is being violated. Note: A nonattainment area might be nonattainment in one
pollutant, but attainment in another. LAER is the most stringent emission limita-
tion which is contained in the SIP of any State or the most stringent emission
limitation achieved in practice by that source category. LAER may be in some
cases considered to be technology forcing, involving the transfer of technology
from one source category to another. LAER is more stringent than BACT or
RACT and must be at least as stringent as the control specified by the NSPS.
1-2
-------
Control equipment which satisfies the requirements of RACT, BACT or LAER is
generally specific to a particular source category. For a given source category one
control device may be able to satisfy all three levels of control, or three different
control devices may be needed. Once the appropriate gaseous emission control
device is identified (absorber, adsorber, or thermal incinerator), its operation
usually can be modified to meet the required level of control. For example, a large
utility may choose to meet the RACT requirements for SO2 emissions by burning
low sulfur coal. BACT or LAER requires scrubbing the flue gas. The same scrub-
ber system (i.e. venturi, spray, or plate) may be used to meet the appropriate con-
trol level by improving its operating efficiency or by adding stages to it.
As illustrated in the above example, gaseous emission control can also be
achieved by methods other than add-on control devices. Changing fuel sources,
modifying or changing raw materials, or using alternative production procedures
are also used to reduce emissions without adding control equipment.
This manual is intended to be used in APTI Course 415 Control of Gaseous
Emissions, and it presents the fundamental concepts of the operation of gaseous
emission control equipment for stationary sources. Control techniques applied to
reducing the emissions of gaseous pollutants are absorption, adsorption, combus-
tion and condensation. Controlling SO2 and NO, emissions from fossil-fuel-fired
steam generators is also discussed. These are the largest stationary sources emitting
SO2 and NO, pollutants. Sources and control of volatile organic compounds are not
included in this manual since this topic is covered in APTI Course 482 Sources and
Control of Volatile Organic Compounds. Also, a chapter describing the auxiliary
equipment, such as hooding, ductwork, and fans associated with the control equip-
ment is included.
Gaseous Emission Control Techniques
Gaseous pollutants are emitted from a variety of processes. Gaseous pollutants
which must be controlled are not limited to those listed in Table 1-1. Most noxious
or toxic gases are also controlled. Table 1-2 lists some typical gaseous pollutants
and their sources (Parekh, 1975).
1-3
-------
Table 1-2. Typical gaseous pollutants and their sources.
Key elemen
S
N
C
Halogen
F
Cl
Pollutant
SO2
SO3
H2SO4 vapors
H2S
R-SH (mercaptans)
NO, NO2
HNOj vapors
NH,
Other N compounds
(i.e. amines, pyridines)
Inorganic
CO
Organics: Volatile Organic
Compounds
Hydrocarbons
Paraffins
Olefins
Aromatics
Oxygenated hydrocarbons
Aldehydes
Ketones
Alcohols
Phenols
Chlorinated solvents
HF
iF4
HC1
C!2
Source
Boiler flue gas
Sulfuric acid manufacturing
Sulfuric acid manufacturing, pickling
operations
Natural gas processing, pulp and paper
mills, sewage treatment
Petroleum refining, pulp and paper mills
Nitric acid manufacturing, boiler flue
gas
Nitric acid manufacturing, pickling
operations
Ammonia manufacturing
Sewage, rendering, solvent processes
Incomplete combustion
Solvent uses, gasoline marketing,
petrochemical plants
Surface coating operations, petroleum
processing, plastics manufacturing
Dry cleaning, degreasing
Phosphate fertilizer plant, aluminum
plant
Ceramics, fertilizer plant
HC1 manufacturing, PVC combustion
Chlorine manufacturing
Techniques used to control gaseous emissions are absorption, adsorption com-
bustion, and condensation. The applicability of a given technique depends on the
physical and chemical properties of the pollutant and the exhaust stream. More
than one technique may be capable of controlling emissions from a given source
For example, vapors generated from loading gasoline into tank trucks at large bulk
terminals are controlled by using any of these four techniques. Most often,
however, one control technique is used more frequently than others for a given
source-pollutant combination. For example, absorption is commonly used to
remove SO2 from boiler flue gas.
1-4
-------
Absorption
Absorption is a mass transfer operation in which a gas is dissolved in a liquid. A
contaminant (pollutant) exhaust stream contacts a liquid, and the contaminant
diffuses from the gas phase into the liquid phase. The absorption rate is enhanced
by high diffusion rates of the contaminant in both the liquid and gas phase, high
solubility of the contaminant, large liquid-gas contact area and, good mixing
between liquid and gas phases (turbulence).
The liquid most often used for absorption is water since it is inexpensive, readily
available, and can dissolve a number of contaminants. Reagents can be added to
the absorbing water to increase the removal efficiency of the system. Certain
reagents merely increase the solubility of the contaminant in the water. Other
reagents chemically react with the contaminant after it has been absorbed (such as
a lime slurry used to scrub SO2 flue gases). In reactive scrubbing the absorption
rate is much higher, so in some cases a smaller, more economical system can be
used. However, the reactions can form a precipitate which could cause plugging
problems in the absorber or associated equipment.
Gas absorbers or wet scrubbers are designed to provide good mixing of the gas
and liquid phases. The devices used for gas absorption are often the same as those
used in particulate emission scrubbing. These include packed towers, plate towers,
spray columns, and venturi scrubbers. Although these devices are identical, a wet
particulate emission scrubber is not operated the same way as a gas absorber is
operated. In a wet scrubber, particulate matter removal efficiency is a function of
the pressure drop-the higher the pressure drop the greater the removal efficiency.
For gas absorption, a higher pressure drop results in shorter contact time between
the phases which limit the amount of absorption that can occur. Therefore,
optimizing both gas and particulate pollutant removal in one device is difficult.
If a gaseous contaminant is very soluble, almost any of the wet scrubbers will
adequately remove this contaminant. However, if the contaminant is of low
solubility, the packed tower or the plate tower is more effective. Both of these
devices provide long contact time between phases and have relatively low pressure
drops. The packed tower, the most common gas absorption device, consists of an
empty shell filled with packing. Liquid flows down over the packing exposing a
large film area to the gas flowing up through the packing. Plate towers consist of
horizontal plates placed inside the tower. Gas passes up through orifices in these
plates, while liquid flows down across the plate providing the desired contact. The
major disadvantage in using these devices is that they both are susceptible to
plugging when the exhaust stream contains high concentrations of particulate
matter. When an exhaust stream contains a high level of particulate matter, ven-
turis, spray devices, or combinations of systems can be used.
Adsorption
Adsorption is a mass transfer process that involves removing a gaseous contaminant
by adhering it to the surface of a solid. Adsorption can be classified as physical or
chemical. In physical adsorption, a gas molecule adheres to the surface of the solid
1-5
-------
due to an imbalance of natural forces (electron distribution). In chemisorption,
once the gas molecule adheres to the surface, it reacts chemically with it. The
major distinction is that physical adsorption is readily reversible while chemisorp-
f"ir»Ti ic rifil- "
tion is not.
All solids physically adsorb gases to some extent. Certain solids, called adsorbents
have a high attraction for specific gases plus a large surface area which gives them a
large capacity. By far the most important adsorbent for air pollution control is
activated carbon. Due to its surface, activated carbon will preferentially adsorb
hydrocarbon vapors and odorous organic compounds from an air stream. Most
other adsorbents (molecular sieves, silica gel, and activated aluminas) will
preferentially adsorb water vapor thereby rendering them useless to remove other
contaminants.
For activated carbon, the amount of hydrocarbon vapors that can be adsorbed
depends on the physical and chemical characteristics of the vapors, their concentra-
tion in the gas stream, system temperature, system pressure, humidity of the gas
stream, and molecular weight of the vapor. Physical adsorption is a reversible
process; the adsorbed vapors can be released by increasing the temperature,
decreasing the pressure, or using a combination of both. Vapors are normally
desorbed by heating the adsorber with steam.
Adsorption can be a very useful removal technique, since it is capable of
removing very small quantities (a few ppm) of vapor from an air stream. The
vapors are not destroyed, only stored on the adsorbent surface until they can be
removed by desorption. The desorbed vapor stream is highly concentrated.
Depending on the nature of this material, it can be condensed and recycled or
burned as the ultimate disposal technique.
The most common adsorption system is the fixed bed adsorber. These systems
consist of two or more adsorber beds operating on a timed adsorbing and desorbing
cycle. One or more beds are adsorbing vapors, while the other bed is being
regenerated. If paniculate matter or liquid droplets are present in the vapor-laden
air stream, this stream is sent to pretreatment to remove them. If the temperature
of the inlet vapor stream is high (much above 120°F), cooling may also be pro-
vided. Also, since all adsorption processes are exothermic, cooling coils in the car-
bon bed itself may be needed to prevent excessive heat buildup. Carbon bed depth
is usually limited to a maximum of four feet and vapor velocity through the
adsorber is held below 100 fpm to prevent an excessive pressure drop.
Combustion
Combustion is defined as rapid, high temperature gas phase oxidation. Simply, the
contaminant (a carbon-hydrogen substance) is burned with air and converted to
carbon dioxide and water vapor. The operation of any combustion source is
governed by the three T's of combustion: temperature, turbulence, and time. For
complete combustion to occur, each contaminant molecule must come in contact
with oxygen (turbulence) at a sufficient temperature, and be held at this
temperature for an adequate time. These three variables are totally dependent on
each other. For example, if a higher temperature were used, less mixing of the
1-6
-------
contaminant and combustion air or shorter residence time may be used. If
adequate turbulence cannot be provided, a higher temperature or longer residence
time may be required for complete combustion.
Combustion devices can be categorized as flares, thermal incinerators, or
catalytic incinerators. Flares are direct combustion devices used to dispose of large
quantities or emergency releases of combustible gases. Flares are normally elevated
(from 100 to 400 ft) to protect the surroundings from heat and flames. Flares are
often designed for steam injection at the flare tip. The steam provides sufficient
turbulence to ensure complete combustion which prevents smoking. Flares are very
noisy which can cause problems for adjacent neighborhoods.
Thermal incinerators are also called afterburners, direct flame incinerators, or
thermal oxidizers. These are devices in which the contaminant air stream passes
around or through a burner and into a refractory-lined residence chamber where
oxidation occurs. To ensure complete combustion of the contaminant, thermal
incinerators are designed to operate between a temperature of 700 to 800 °C (1300
to 1500°F) and a residence time of 0.3 to 0.5 seconds. Ideally, as much of the fuel
value as possible is supplied by the waste contaminant stream. This reduces the
amount of auxiliary fuel needed to maintain the proper temperature.
In catalytic incineration the contaminant-laden stream is heated and passed
through a catalyst bed which promotes the oxidation reaction at a lower
temperature. Catalytic incinerators normally operate between 370 and 480 °C (700
and 900°F). This reduced temperature represents a continuous fuel savings.
However, this may be offset by the cost of the catalyst. The catalyst, which is
usually platinum, is coated on a cheaper metal or ceramic support base. The sup-
port can then be arranged to expose a high surface area, which provides sufficient
active sites on which the reactions occur. Catalysts are subject to both physical and
chemical deterioration. Halogens and sulfur containing compounds act as catalyst
suppressants and decrease the catalysts' usefulness. Certain heavy metals such as
mercury, arsenic, phosphorous, lead and zinc are particularly poisonous.
Condensation
Condensation is a process in which the volatile gases are removed from the con-
taminant stream and changed to a liquid. Condensation is usually achieved by
reducing the temperature of a vapor mixture until the partial pressure of the con-
densable component equals its vapor pressure. Condensation requires low
temperatures to liquify most pure contaminant vapors. Condensation is affected by
the composition of the contaminant gas stream. The presence of additional gases
which do not condense at the same conditions, such as air, hinders condensation.
Condensers are normally used in combination with primary control devices. Con-
densers can be located upstream of (before) an incinerator, adsorber or absorber.
These condensers reduce the volume of vapors the more expensive equipment must
handle. Therefore, the size and the cost of the primary control device can be
reduced. Similarly, condensers can be used to remove water vapors from a process
stream with a high moisture content upstream of a control system. A prime exam-
ple is the use of condensers in rendering plants to remove moisture from the cooker
1-7
-------
exhaust gas. When used alone, refrigeration is required to achieve the low
temperatures required for condensation. Refrigeration units are used to successfully
control gasoline vapors at large bulk gasoline dispensing terminals.
Condensers are classified as being either contact condensers or surface con-
densers. Contact condensers cool the vapor by spraying liquid directly on the vapor
stream. These devices resemble a simple spray scrubber. Surface condensers are
normally shell-and-tube heat exchangers. Coolant flows through the tubes, while
the vapor is passed over and condenses on the outside of the tubes. In general, con-
tact condensers are more flexible, simpler and less expensive than surface con-
densers. However, surface condensers require much less water and produce nearly
20 times less wastewater that must be treated than contact condensers. Surface con-
densers also have an advantage in that they can directly recover valuable contami-
nant vapors.
This chapter presented a brief overview of the regulations which specify a
required level of emission reduction. The four methods used to control gaseous
emissions were summarized. These methods will be discussed in detail in later
chapters in this manual. The chapters on the control methods have been developed
in three steps. First, the basic theory of the control method is presented, then a sec-
tion discusses procedures used to estimate size and operating conditions of the con-
trol equipment, and finally, the types and operation of the equipment are
described. The equations presented in each chapter are used to estimate important
design variables for that particular control device. They are intended to give prac-
tical operating values, not rigid design values. The following chapter reviews some
of the basic concepts of the behavior of gases. These basic concepts are important
for understanding the operation of gaseous emission control equipment.
Reference
Parekh, R. 1975. Equipment for controlling gaseous pollutants. Chem. Engr.
82:129 (October 6).
1-8
-------
Chapter 2
Basic Concepts of Cases
In order to properly evaluate gaseous control devices, a thorough understanding of
the process variables that affect a gas stream is essential. This chapter briefly
reviews a few basic concepts of gas behavior. The chapter is divided into two parts.
The first half defines important physical and chemical properties of gases:
temperature, pressure, density, molecular weight, viscosity and specific heat. Com-
mon symbols, units of measurement, and any important values are presented. The
second half of this chapter reviews a few scientific laws used to predict the behavior
of a gas under varying process conditions including the ideal gas law, Raoult's Law,
Dalton's Law and Henry's Law.
The system of measurement used in this manual is the International System of
Units (SI). This system is an updated, more restrictive version of the metric system
which derives its base units from physical phenomena and scientific relationships.
In this manual the SI units are normally listed first with the English units given in
parenthesis. Occasionally the units in parenthesis are the commonly used units
instead of English units. For some tables, figures, and equations, conversion from
English to SI was not practical. For these instances only the English units are
presented.
Expression of Gas Temperature
Fahrenheit and Celsius Scales
The range of units on the Fahrenheit scale between freezing and boiling is 180; on
the Celsius or Centigrade scale, the range is 100. Therefore, each Celsius-degree is
equal to 9/5 or 1.8 Fahrenheit-degree. The following relationships convert one
scale to the other:
(Eq. 2-1) °F=1.8°C + S2
(Eq. 2-2) °C = (°F - 32)/1.8
Where: °F = degrees Fahrenheit
°C = degrees Celsius or degrees Centigrade
Absolute Temperature
Experiments with perfect gases have shown that, under constant pressure, for each
change in Fahrenheit-degree below 32°F the volume of gas changes 1/491.67.
Similarly, for each Celsius-degree, the volume changes 1/273.16. Therefore, if this
change in volume per temperature-degree is constant, the volume of gas would,
theoretically, become zero at 491.67 Fahrenheit-degrees below 32°F or at a reading
2-1
-------
of -459.67°F. On the Celsius or Centigrade scale, this condition occurs at 273.16
Celsius-degrees below 0°C, or at a temperature of -273.16°C. Relationships of the
various temperatures are shown in Figure 2-1.
Fahrenheit
212
-459.6
Absolute
(K)
671.6 1 I 373
491.6
273
Celsius
100
-273
Figure 2-1. Comparison of Fahrenheit, Celsius, and absolute temperature scales.
Absolute temperatures determined by using Fahrenheit units are expressed as
degrees Rankine (°R); those determined by using Celsius units are expressed as
degrees Kelvin (K). The following relationships convert one scale to the other:
°R= °F + 459.67
K= °C + 273.16
(Eq. 2-3)
(Eq. 2-4)
In the International System of Units (SI) the degree mark (°) is not used with the
units in the Kelvin scale. Throughout this manual degrees Celsius will be °C,
degrees Kelvin will be K.
Expression of Gas Pressure
Definition of Pressure
A body may be subjected to three kinds of stress: shear, compression, and tension.
Fluids are unable to withstand tensile stress; hence, they are subject to shear and
compression only.
Unit compressive stress in a fluid is termed pressure and is expressed as force per
unit area (e.g. N/m2 or Newton/meter2, lb/in2 or psi). For the SI system of units,
1 Newton/meter2 is defined as a Pascal (Pa).
Pressure is equal in all directions at a point within a volume of fluid, and acts
perpendicular to a surface.
2-2
-------
Barometric Pressure
Barometric pressure and atmospheric pressure are synonymous. These pressures are
measured with a barometer and are usually expressed as inches, or millimeters, of
mercury (Hg). Standard barometric pressure is the average atmospheric pressure at
sea level, 45° north latitude at 35 °F. It is equivalent to a pressure of 14.696
pound-force per square inch exerted at the base of a column of mercury
29.92 inches high. In the SI system this pressure is equal to 101.325 kPa. In the
metric system, the average atmospheric pressure is 760 mm Hg or 1 atm. Weather
and altitude are responsible for barometric pressure variations.
Gage Pressure
Measurements of pressure by ordinary gages are indications of difference in
pressure above, or below, that of the atmosphere surrounding the gage. Gage
pressure, then, is ordinarily the pressure of the system. If greater than the pressure
prevailing in the atmosphere, the gage pressure is expressed as a positive value; if
smaller, the gage pressure is expressed as negative. The term vacuum designates a
negative gage pressure.
The abbrevation g (in pg) is used to specify a gage pressure. For example, psig
means pound-force per square inch gage pressure.
Absolute Pressure
Because gage pressure (which may be either positive or negative) is the pressure
relative to the prevailing atmospheric pressure, the gage pressure, added
algebraically to the prevailing atmospheric pressure (which is always positive), pro-
vides a value that has a datum of absolute zero pressure. A pressure calculated in
this manner is called absolute pressure. The mathematical expression is:
(Eq. 2-5) P=P* + p,
Where: P= absolute pressure, kPa
Pb = barometric pressure (atmospheric), kPa
pg —gage pressure, kPa
The abbrevation a is sometimes used to indicate that the pressure is absolute.
For example, psia means pounds per square inch absolute pressure. The symbol P
by itself, will also be used in this manual to indicate absolute pressure.
2-3
-------
The absolute pressure allows conversion of one pressure system to the other.
Relationship of the pressure systems are shown graphically in Figure 2-2 using two
typical gage pressures, p4, and pgt. Gage pressure pgl is above the zero from which
gage pressures are measured, and, hence, is expressed as a positive value; gage
pressure pg2 is below the gage pressure zero, and, therefore, is expressed as a
negative value.
p,i (Positive valve)
P«i (Negative valve)
p,, zero
P, zero
Figure 2-2. Gas-pressure relationship.
Density
Density is defined as a mass per unit volume:
(Eq. 2-6) Q = —
Where: Q = density, kg/m3
M = mass of a substance, kg
V = volume, m3
The density of a solid or a liquid is a function of the temperature at which the
density is measured. For example, water has a density of 1.0 g/cm3 (62.42 lb/ft3)
at 0°C. Gas densities are referenced at a given temperature and pressure. For
example, the density of dry air at 101.325 kPa (760 mm Hg) and 25 °C is 1.1844
mg/cm3 (0.0739 lb/ft3). Densities of specific compounds may be found in tables
listed in handbooks such as Perry's Chemical Engineering Handbook.
Appendices G and H give some of the physical properties of air and water.
2-4
-------
Molal Units
The molecular weight of a compound or an element is simply the sum of all the
atomic weights (from the Periodic Chart) of all the atoms in the molecule. A mole
of any pure substance is defined as the amount of substance that numerically
equals the molecular weight of that compound. Therefore, by definition, a gram
mole or pound mole of oxygen weighs 32 grams or pounds depending on the units
used. Molal units are used extensively in air pollution control for gaseous removal
calculations as they greatly simplify material balances where chemical reactions
(removing pollutants) are occurring.
For mixtures of substances (gas, liquid and solids) it is also convenient to express
compositions in mole fraction or mole percent instead of a mass fraction. The mole
fraction is the ratio of the number of moles of one component to the total number
of moles in the mixture. For control of gaseous pollutants, these mixtures involve
an exhaust stream (and sometimes a liquid stream) with a certain concentration of
pollutant.
Viscosity
Origin and Definition of Viscosity
Viscosity is associated with the fluid resistance to flow. Viscosity is the result of two
physical phenomena: intermolecular cohesive forces, and momentum transfer
between flowing strata caused by molecular agitation perpendicular to the direction
of motion. Between adjacent strata of a moving fluid, a shearing stress (T) that is
directly proportional to the velocity gradient occurs (Figure 2-3).
Figure 2-3. Shearing stress in a moving fluid.
2-5
-------
This is expressed in the equation:
(Eq. 2-7) T = M —
dy
Where: T = unit shearing stress between adjacent layers of fluid
dv/dy= velocity gradient, m/s
/i = viscosity proportionality constant, centipoise
The proportionality constant, p, is called the coefficient of viscosity, or merely,
viscosity. It should be noted that the pressure does not appear in Equation 2-7
which indicates that the shear (T) and the viscosity (/t) are independent of pressure.
(Viscosity actually increases very slightly with pressure but this variation is neg-
ligible in most engineering problems.)
Kinematic Viscosity
The ratio of the absolute viscosity to the density of a fluid often appears in dimen-
sionless numbers such as the Reynolds Number. The expression for kinematic
viscosity is used to simplify calculations. Kinematic viscosity is defined as:
(Eq. 2-8) «*=-£-
e
Where: v — kinematic viscosity, mVs
/«i = viscosity of the gas, Pa»s
Q = density of the gas, g/cm3
Liquid Viscosity
In a liquid, transfer of momentum between strata having a relative velocity is small
compared to the cohesive forces between molecules. Hence, shear stress is
predominantly the result of intermolecular cohesion. Because forces of cohesion
decrease rapidly with an increase in temperature, the shear stress decreases with an
increase in temperature. Equation 2-7 shows that shear stress is directly propor-
tional to the viscosity. Therefore, liquid viscosity decreases when the temperature
increases.
Gay Viscosity
In a gas, the molecules are too far apart for intermolecular cohesion to be effec-
tive. Thus, shear stress is predominantly the result of an exchange of momentum
between flowing strata caused by molecular activity. Because molecular activity
increases as temperature increases, the shear stress increases with a rise in the
temperature. Therefore, gas viscosity increases as the temperature rises.
2-6
-------
Determination of Viscosity of Gases
The viscosity of a gas for the given conditions may be found accurately from the
following formula:
(Eq. 2-9)
273.2
Where: n = viscosity of the gas for prevailing conditions, Pa»s
/x° = viscosity at 0°C and prevailing pressure, Pa«s
T= absolute prevailing temperature, K
n = an empirical exponent (n= 0.768 for air)
The viscosity of air a ad other gases at various temperatures and at a pressure of
1 atmosphere may be found in engineering tables. One unit commonly used to
describe viscosity has been the centipoise which equals 0.01 g/cm»s (6.72x 10"*
lb/ft»s). In the SI system this now becomes a millipascal second (1 cp= 1 mPa»s).
Specific Heat
The specific heat of a gas is the amount of heat required to change the
temperature of a unit-mass of gas one temperature-degree. Units of specific heat
are calories/g» °C or Btu/lb»°F, depending upon the dimensional system used.
Heat may be added while the volume or pressure of the gas remains constant.
Hence, there may be two values of specific heat; the specific heat at constant
volume (C,,) and the specific heat at constant pressure (Cp).
Because the heat energy added at constant pressure is used in raising the
temperature and doing work against the pressure as expansion takes place, Cp is
greater than Cv. The most commonly encountered heat capacity is Cp and unless
otherwise noted will be the value used in this manual when referring to specific
heat. Since Cp varies with temperature, an average specific heat between two
temperature values is often used. The average specific heat will be denoted as Cp.
Values for the specific heats of common gases and liquids are given in Appendix F
of this manual.
Relationships of Ideal Gases
Laws of Boyle and Charles
Boyle's Law states that, when the temperature (T) is held constant, the volume (V)
of a given mass of a perfect gas of a given composition varies inversely as the
absolute pressure varies, i.e.:
V« —at constant T
P
Where: oc = proportional to
2-7
-------
Charles' Law states that, when the volume is held constant, the absolute
pressure of a given mass of a perfect gas of a given composition varies directly as
the absolute temperature varies, i.e.:
—at constant V
T
Where:
-------
The gas law applies to mixtures of gases as well as a pure gas. Thus a gram mole
of ideal gas, or mixture of ideal gases, occupies 22.41 liters at 0°C and 101.3 kPa.
This volume is called the gram-molal volume. Similarly, a pound mole of ideal gas
occupies 359 ft3 at 32 °F and 1 atm. The pound-molal volume at any other com-
bination of temperature and pressure may be calculated by Equation 2-11.
(Eq. 2-11) V2 = V, ±
Pt T,
Dal ton's Law of Partial Pressure
When gases, or vapors (having no chemical interaction), are present as a mixture
in a given space, the pressure exerted by a component of the gas mixture at a given
temperature is the same as it would exert if it filled the whole space alone. The
pressure exerted by one component of a gas mixture is called its partial pressure.
The total pressure of the gas mixture is the sum of the partial pressures. For com-
ponent A, Dalton's Law is expressed as:
9 19^ « _ MiRTV
2-12) -~~
Where: pA = partial pressure of component air mixture, Pa
MX = mass of A, g
MWX = molecular weight of A, g mole
V = volume of the mixture, ms
T = absolute prevailing temperature, K
A useful concept from this and the ideal gas law is pressure percent = volume
percent = mole percent.
For example, if a mixture of air is 21% oxygen and 79% nitrogen by volume at
101.3 kPa (1 atm) and 0°C, then this mixture contains 0.21 mole fraction oxygen
and 0.79 mole fraction nitrogen. Also, the partial pressure of oxygen is equal to
21.27 kPa (0.21 atm) and the partial pressure of nitrogen is 80.03 kPa (0.79 atm)
at 0°C. In mathematical terms the mole fraction in the gas phase is related to the
partial pressure by:
(Eq. 2-13) yA= El
— i
Where: px = partial pressure of A, Pa
PT= total pressure of the system, Pa
2-9
-------
Gas-Liquid Relationships
Raoult's Law states that the partial pressure of a component over a solution is the
product of the vapor pressure of that component and the mole fraction of that
component. Raoult's Law is expressed as:
-------
The larger the Reynolds Number, the smaller is the effect of viscous forces; the
smaller the Reynolds Number, the greater the effect of the viscous forces.
The linear dimension, /, is a length characteristic of the flow system. It is equal
to four times the mean hydraulic radius, which is the cross sectional area divided
by the wetted perimeter. Thus, for a circular pipe, / = diameter of the pipe; for a
particle settling in a fluid medium, 1= diameter of the particle; for a rectangular
duct, /= twice the length times the width divided by the sum; and for an annulus
such as a rotameter system, / = outer diameter minus the inner diameter.
Laminar and Turbulent Flow
In laminar flow, the fluid is constrained to motion in layers (or laminae) by the
action of viscosity. The layers of fluid move in parallel paths that remain distinct
from one another; any agitation is of a molecular nature only. Laminar flow occurs
when Reynolds Number for circular pipes is less than 2000 and less than 0.1 for
particles settling in a fluid medium.
In turbulent flow, the fluid is not restricted to parallel paths but moves forward
in a haphazard manner. Fully turbulent flow occurs when the Reynolds Number is
greater than 2500 for circular pipes and greater than 1000 for settling particles.
2-11
-------
Chapter 3
Combustion
Introduction
The process of combustion is most often used to control the emissions of volatile
organic compounds from process industries. At a sufficiently high temperature and
adequate residence time, any organic vapor can be oxidized to carbon dioxide and
water by the combustion process. Combustion systems are often relatively simple
devices capable of achieving very high removal efficiencies. They consist of:
burners, which ignite the fuel and organic vapors; and a chamber, which provides
appropriate residence time for the oxidation process. Due to the high cost and
decreasing supply of fuels, combustion systems are designed to include some type of
heat recovery. If heat recovery can be used, combustion can be a very effective
control technique. For example, pollutant emissions from paint bake ovens can be
reduced by 99.9+ % using incineration while heat recovered from the incinerator
flue gases can be fed back to the oven. Combustion can also be used for serious
emission problems which require high destruction efficiencies, such as odor
problems or the emission of toxic gases.
There are, however, some problems that may occur when using combustion to
control gaseous pollutants. Incomplete combustion of many organic compounds
results in the formation of aldehydes and organic acids which may create an addi-
tional pollution problem. Oxidizing organic compounds containing sulfur or
halogens produce unwanted pollutants such as sulfur dioxide, hydrochloric acid,
hydrofluoric acid or phosgene. If present, these pollutants would require a scrubber
to remove them prior to release into the atmosphere.
Four basic combustion systems can be used to control combustible gaseous emis-
sions. They are flares, thermal oxidizers, catalytic oxidizers, and process boilers.
Although these devices are physically similar, the parameters under which they
operate are markedly different. Choosing the proper device depends on many fac-
tors including: concentration of combustibles in the gas stream, process flow rate,
control requirements, presence of contaminants in the waste stream, and an
economic evaluation. This chapter will examine the principles of combustion, sim-
ple combustion calculations, and the design and operating parameters of each
combustion device.
Combustion Principles
Combustion is a chemical process occurring from the rapid combination of oxygen
with various elements or chemical compounds, resulting in the release of heat. The
process of combustion is also referred to as oxidation. Most fuels used for combus-
3-1
-------
tion are composed essentially of carbon and hydrogen, but can include other
elements such as sulfur. Simplified reactions for the oxidation of carbon and
hydrogen are given as:
-------
The ignition temperature of various fuels and compounds can be found in com-
bustion handbooks such as the North American Combustion Handbook (1965).
These temperatures are dependent on combustion conditions and therefore should
be used only as a guide. Ignition depends on (EPA, 1972):
1. concentration of combustibles in the waste stream
2. inlet temperature of the waste stream
3. rate of heat loss from combustion chamber
4. residence time and flow pattern of the waste stream
5. combustion chamber geometry and materials of construction.
Most incinerators operate at higher temperatures than the ignition temperature
which is a minimum temperature. Thermal destruction of most organic compounds
occurs between 590 and 650°C (1100 and 1200°F). However, most incinerators are
operated at 700 to 820 °C (1300 to 1500°F) to convert CO to CO2; which occurs
only at these higher temperatures.
Time
Time and temperature affect combustion in much the same manner as
temperature and pressure affect the volume of a gas. When one variable is
increased, the other may be decreased with the same end result. With a higher
temperature, a shorter residence time can achieve the same degree of oxidation.
The reverse is also true, a higher residence time allows the use of a lower
temperature. In describing incinerator operation, these two terms are always men-
tioned together. One has little meaning without specifying the other.
The choice between higher temperature or longer residence time is based on
economic considerations. Increasing residence time involves using a larger combus-
tion chamber resulting in a higher capital cost. Raising the operating temperature
increases fuel usage which also adds to the operating costs. Fuel costs are the major
operating expense for most incinerators. Within certain limits, lowering the
temperature and adding volume to increase residence time can be a cost effective
alternative method of operation.
The residence time of gases in the combustion chamber may be calculated from:
(Eq. 3-3) 0SS"o~
Where: 0 = residence time, s
V = chamber volume, m3
Q= gas volumetric flow rate at combustion conditions, mVs
Q is the total flow of hot gases in the combustion chamber. Adjustments to the
flow rate must include any outside air added for combustion. Example 3-1 shows
the determination of residence time from the volumetric flow rate of gases.
3-3
-------
Turbulence
Proper mixing is important in combustion processes for two reasons. First, for com-
plete combustion to occur, every particle of fuel must come in contact with air
(oxygen). If not, unreacted fuel will be exhausted from the stack. Second, not all
of the fuel or waste gas stream is able to be in direct contact with the burner
flame. In most incinerators a portion of the waste stream bypasses the flame
and is mixed at some point downstream of the burner with the hot products of
combustion. If the two streams are not completely mixed, a portion of the waste
stream will not react at the required temperature and incomplete combustion will
occur.
A number of methods are available to improve mixing the air and combustion
streams. Some of these include the use of refractory baffles, swirl fired burners,
and baffle plates. These devices are discussed in more detail in the equipment sec-
tion of this chapter. The problem of obtaining complete mixing is not easily solved.
Unless properly designed, many of these mixing devices may create "dead spots"
and reduce operating temperatures. Merely inserting obstructions to increase tur-
bulence is not the answer. According to one study of afterburner systems, "the
process of mixing flame and fume stream to obtain a uniform temperature for
decomposition of pollutants is the most difficult part in the design of the after-
burner" (EPA, 1972).
Oxygen Requirements
Oxygen is necessary for combustion to occur. To achieve complete combustion of a
compound, a sufficient supply of oxygen must be present to convert all of the car-
bon to CO2. This quantity of oxygen is referred to as the stoichiometric or
theoretical amount. The stoichiometric amount of oxygen is determined from a
balanced chemical equation summarizing the oxidation reactions. For example,
from Equation 3-4, 1 mole of methane requires 2 moles of oxygen for complete
combustion.
(Eq- 3-4) CH4 + 202-C02 + 2H2O
If an insufficient amount of oxygen is supplied, the mixture is referred to as rich.
There is not enough oxygen to combine with all the fuel so that incomplete com-
bustion occurs. This condition results in black smoke being exhausted. If more
than the stoichiometric amount of oxygen is supplied, the mixture is referred to as
lean. The added oxygen plays no part in the oxidation reaction and passes through
the incinerator.
Oxygen for combustion processes is supplied by using air. Since air is essentially
79% nitrogen and 21% oxygen, a larger volume of air is required than if pure
oxygen were used. To balance Equation 3-4, 9.53 moles of air would be required
to completely combust the 1 mole of methane. The stoichiometric calculations for
this are presented in Appendix D. A listing of theoretical air requirements is given
in Table 3-1.
3-4
------- An error occurred while trying to OCR this image.
-------
In industrial applications, more than the stoichiometric amount of air is used to
ensure complete combustion. This extra volume is referred to as excess air. If ideal
mixing were achievable, no excess air would be necessary. However, most combus-
tion devices are not capable of achieving ideal mixing of the fuel and air streams.
The amount of excess air is held to a minimum in order to reduce heat losses.
Excess air takes no part in the reaction but does absorb some of the heat produced.
To raise the excess air to the combustion temperature, additional fuel must be used
to make up for this loss of heat. Operating at a high volume of excess air can be
very costly in terms of the added fuel required. Equations to calculate excess air
are also listed in Appendix D.
In addition to the theoretical air required, Table 3-1 also lists the volume of
combustion products produced from oxidizing a substance. This is an important
term used to determine the size of the combustion chamber. Example 3-1 illustrates
how these values are used. The values in Table 3-1 are given in volume percent
and weight percent. Volume percent is the more important term for combustion of
gaseous substances since gas flows are measured in cubic meters per second instead
of weight units. For example, from Table 3-1, when 1 m3 of methane is combusted
with the theoretical amount of air, 10.53 m3 of flue gas is produced. Table 3-1 is
given in English units. Since these are volume ratios (ftVft3), metric units and
English units are interchangeable.
Natural gas is not listed in Table 3-1 since its chemical composition can vary.
When 1 m3 of natural gas is burned with a stoichiometric amount of air, it pro-
duces approximately 11.5 m3 (average value) of flue gas.
Combustion Limits
Not all mixtures of fuel and air are able to support combustion. The flammable or
explosive limits for a mixture are the maximum and minimum concentrations of
fuel in air that will support combustion. The upper explosive limit (UEL) is
defined as the concentration of fuel which produces a nonburning mixture due to a
lack of oxygen. The lower explosive limit (LEL) is defined as the concentration of
fuel below which combustion will not be self-sustaining. Table 3-1 lists the
flammability limits (LEL and UEL) for common fuels and solvents.
For example, consider that a mixture of gasoline vapors and air is at
atmospheric conditions. From Table 3-1 the LEL is 1.4% by volume of gasoline
vapors and the UEL is 7.6%. Any concentration of gasoline in air within these
limits will support combustion. That is, once a flame has been ignited it will con-
tinue to burn. Concentrations of gasoline in air below or above these limits will not
burn and can quench the flame. The lower explosive limit is the more important of
the two terms in describing gas streams with combustible contaminants. Industrial
processes which handle combustible vapors, such as paint or solvent vapors, are
usually required by insurance companies to operate at less than 25% of the LEL of
the vapor in the ducts to minimize fire hazards. By using gas analyzers and an
alarm system, the concentration of vapor may be allowed to be as high as 50% of
the LEL by an insurance company covering the plant.
Flammability limits are not absolute values, but are affected by temperature,
3-6
-------
pressure, geometry of the chamber, and presence of other contaminants. The
higher the temperature, the greater the activation energy of combustion is. Thus,
there is a greater probability that the flame will propagate. At high enough
temperatures even vapor concentrations within 25% of their LEL can propagate a
flame. This high degree of flammability limits the temperature to which some
waste streams can be preheated prior to oxidation. Pressure also changes vapor
flammability by its effect on gas densities. The higher the pressure, the closer the
gas molecules are to one another. This decreases the distance a flame or spark
must jump from one combustible point to another (Bethea, 1978).
The size and geometry of the combustion chamber also affect the flammability
limits. The smaller the diameter of the chamber, the narrower the flammability
range will become. Increased surface-to-volume ratio promotes rapid cooling and
flame quenching, limiting flammability. Some flame arresters are designed on this
principle of large surface-to-air volume ratio to eliminate fire hazards. The
presence of other compounds can also affect flammability limits. The flammability
range of a mixture can be either increased or decreased depending on the proper-
ties of the contaminant.
Flame Combustion
When mixing fuel and air, two different mechanisms of combustion can occur. A
luminous (yellow) flame results when air and fuel flowing through separate ports
are ignited at the burner nozzle. The yellow flame results from thermal cracking of
the fuel. Cracking occurs when hydrocarbons are intensely heated before they have
a chance to combine with oxygen. The cracking releases both hydrogen and carbon
which diffuse to the flame to form CO2 and H2O. The carbon particles give the
flame the yellow appearance. If incomplete combustion occurs from flame tempera-
ture cooling or if there is insufficient oxygen, soot and black smoke will form.
Blue flame combustion occurs when the fuel and air are premixed in front of the
burner nozzle. This produces a short, intense, blue flame. The reason for the dif-
ferent flame is that the fuel-air mixture is gradually heated. The hydrocarbon
molecules are slowly oxidized, going from aldehydes and ketones to CO2 and H2O.
No cracking occurs and no carbon particles are formed. Incomplete combustion
results in the release of the intermediate, partially oxidized compounds. Blue haze
and odors are emitted from the stack.
Example 3-1
Emissions from a paint baking oven are controlled by a thermal oxidizer. The
unit has a diameter of 1.5 m (5 ft) and is 3.5 m (12 ft) long. The exhaust from
the oven is 3.8 mVs (800 scfm). The oxidizer uses 0.14 mVs (300 scfm) of natural
gas and operates at a temperature of 760°C (1400°F). If all the oxygen necessary
for combustion is supplied from the process stream (no outside air added), what is
the residence time of the gases in the chamber?
Solution:
To solve the problem one can use the approximation given previously that 11.5m3 of
combustion products are formed for every 1.0 m3 of natural gas burned at stan-
3-7
-------
dard conditions (16 °C and 101.3 kPa). Also, 10.33 m3 of theoretical air is required
to combust 1 m3 of natural gas at standard conditions.
1 . First determine the volume of combustion products from burning the natural
gas.
0.14 — x 11-5 m3 of Product = ! 61 E!
s 1.0 m3 of gas s
2. Then determine the air required for combustion.
10.33m3 of air
=1.45 —
1.0 m3 of gas s
Sum up the volumes.
flow from paint bake oven = 3.8 mVs
products of combustion = 1.61 mVs
minus the air from exhaust used in combustion = (- 1.45 mVs)
total volume = 3.96 mVs
Convert to actual mVs.
3.96 x
s (273°C)
3. The volume of the chamber is given by:
= 7r(0.75m)2(3.5m)
= 6.18 m3
4. The residence time from Equation 3-3 is:
Q 6.18m3
0 = - _ o.4
14.98 mVs
Combustion Calculations
Definitions
In describing any combustion process, there are numerous terms used to define
heat. These terms, describing the heat of a system, can be broken into two
categories: combustion and thermodynamic. Combustion terms were initiated to
aid in standardizing fuel usage calculations. These terms are applied to heat that is
produced by combustion methods. Thermodynamic terms apply to all systems.
Thermodynamic terms define the energy level or potential (in terms of heat) that is
present in any substance. Since the combustion terms are specific examples of the
thermodynamic terms, some overlap is involved in defining them.
3-8
-------
The following are important terms describing heat thermodynamically:
Sensible heat (H,): Heat which when added or removed causes a change in
temperature.
Latent heat (Hv): Heat given off by a vapor condensing to a liquid or gained by a
liquid evaporating to a vapor, without a change in temperature. The latent heat of
vaporization of water at 212°F is 970.3 Btu/lb.
Heat content or enthalpy (H): The sum total of latent and sensible heat present
in a substance (gas, liquid or solid) minus that contained at an arbitrary set of
conditions chosen as the base or zero point. Values for various gases are listed in
Table 3-2.
Table 3-2. Heat contents of various gases.
Temp
(°F)
60
100
200
300
400
500
600
700
800
900
1000
1200
1400
1600
1800
2000
2200
2400
2600
2800
3000
3200
3400
3600
Heat content, H
(Btu/lb at 1 atm)
0,
0
8.8
30.9
53.3
76.2
99.4
123.1
147.2
171.7
196.6
221.7
272.5
324.3
377. 3
430.7
484.0
539.3
594.4
649.0
702.8
758.6
816.4
873.4
931.0
N,
0
9.9
34.8
59.9
85.0
110.3
136.1
161.7
187.7
213.9
240.7
294.7
350.8
407.3
465.0
523.8
583.2
642.3
702.8
763.1
824 1
885 8
947.6
1010.3
Air
0
9.6
33.6
57.7
81.8
106.0
130.2
154 5
178.9
203.4
235.0
288.5
343.0
398.0
455.0
51S.O
570.7
628.5
687.3
746 6
806.3
866.0
925.9
986 1
GO
0
10.0
34 9
59.9
85 0
110.6
136.3
162.4
188.7
215.6
242.7
297.8
354.3
407.5
465.3
523.8
583.3
643.0
703.2
771.3
832.6
894.0
956.0
1018 3
CO,
0
8.0
29.3
52.0
75.3
99.8
125 1
149.6
177 8
205.6
233.6
290.9
349 7
416.3
470 9
532.8
596.1
659 2
723.2
787.4
852.0
916 7
981.6
1047.3
SO,
0
5.9
21 4
37 5
54.4
71.8
89.8
108.2
127.0
146 1
165.5
205.1
245.4
286.4
327 8
369.1
411 1
452.7
495.2
557 5
580.0
622.5
665.0
707 5
H,
0
137
484
832
1182
1532
1882
2233
2584
2935
3291
4007
4729
5460
6198
6952
7717
8490
9272
10060
10870
11680
12510
13330
CH.
0
21.0
76.1
136.4
202.1
272.6
347.8
427.4
511.2
599.2
691 1
886.2
1094.1
1313.0
1542.6
-
—
—
-
_
—
—
H,0
0
~
—
1165
1212
1259
1307
1355
1404
1454
1505
1609
1717
1829
—
—
—
—
-
—
—
—
Source: North American Combustion Handbook, North American Manufacturing
Co., Cleveland, OH, 1st ed. (1952).
Some useful terms describing the heat produced by the combustion of a fuel are:
Gross heating value (HVC): The total heat obtained from the complete combus-
tion of a fuel which is at 60 °F when combustion starts, and the combustion
products of which are cooled to 60 °F before the quantity of heat released is
measured. Constant pressure, normally 101.3 kPa (1 atmosphere) is maintained
throughout the entire combustion process. Gross heating values are also referred to
as total or higher heating values.
3-9
-------
Net heating value (H\V): The gross heating value minus the latent heat of
vaporization of the water formed by the combustion of the hydrogen in the fuel.
For a fuel containing no hydrogen, the net and gross heating values are the same.
Available heat (HA): The gross quantity of heat released within a combustion
chamber minus (1) the sensible heat carried away by the dry flue gases and, (2) the
latent heat and sensible heat carried away in water vapor contained in the flue
gases. The available heat represents the net quantity of heat remaining for useful
heating. Figure 3-1 shows the available heat from the complete combustion (no
excess air) of various fuels at various flue gas temperatures.
••5 140,000
120,000
w- 100,000
«
-------
Since all of the previous terms describe heat, they all are expressed in units of
Joules per kilogram (1 Btu/lb is equal to 2324 J/kg). Figure 3-2 illustrates the
interrelation of these terms.
Thermodynamic heat
terms
Heat content (H)
Sensible
heat (H,)
Latent heat
Combustion heat
terms
Gross heating valve (HVC)
Available
heat (HA)
Net heat
Heat lost in
flue gases
Latent heat
of vaporization
(Hv)
Figure 3-2. Heat terms.
Depending on the user, the above terms can also have more than one definition.
For example, a laboratory chemist may describe latent heat as the energy used in
the chemical combustion of a fuel to carbon dioxide and water; while a boiler
operator may describe latent heat as the difference between the gross and net
heating values.
Another important term used in performing combustion calculations is the
specific heat, Cp, of a substance. Specific heat is defined as the amount of heat
required to raise 1 gram of a substance 1 degree centigrade. Specific heat is given
as J/kg»K in SI units and is Btu/lb«°F in English units. Specific heat depends on
temperature. Appendix F gives a nomograph that can be used to compute the
specific heats of common gases at various temperatures.
3-11
-------
Heat Balance
A main area of concern in vapor incineration deals with the amount of fuel
required to raise the temperature of the waste stream to the temperature required
for complete oxidation. The first step in computing the heat required is to perform
a heat balance around the oxidation system. Figure 3-3 shows a typical incineration
system illustrating the heat entering and leaving the incinerator.
Heat losses
Waste gas
TI, ihi
Air and fuel-
T2, m,
Flue gas
• IIIIIIIIIIIMII Illlllllllllllllllll IMIIIIIIIIIIIII
E System boundary -^-
Where: TI = waste gas inlet temperature
Tj = incinerator temperature
rhi = mass flow rate of waste gas plus air and fuel
rh2 = mass flow rate of flue gases
Figure 3-3. Heat balance.
From the law of conservation of energy:
(Eq. 3-5) Heat in = Heat out + Heat lost
Heat is a relative term which is compared at a reference temperature. The heat
content of a substance is arbitrarily taken as zero at a specified reference
temperature. In the gas industry (natural gas), the reference temperature is nor-
mally 16°C (60 °F). This is the reference temperature used throughout this chapter.
Thus, the heat content can be computed from Equation 3-6 as well as found by
using tables such as Table 3-2.
(Eq. 3-6)
H=CP(T-T0)
Where: H = enthalpy, J/kg
Cp = specific heat at temperature T, J/kg.°C
T = temperature of the substance, °C
T0 = reference temperature, °C
3-12
-------
Subtracting the heat going in from the heat leaving the system in Figure 3-3,
gives the heat which must be supplied by the fuel. This is referred to as a change
in enthalpy or heat content. Using Equation 3-6, the enthalpy going in (T,) is sub-
tracted from the enthalpy leaving (T2) giving:
(Eq. 3-7) AH = [Cp2(T2 - T0) - Cpl(T, - T.)]
Where: AH = change in enthalpy
To simplify this calculation, an average specific heat value (Cp) can be used over
the temperature range involved. This reduces Equation 3-7 to:
(Eq. 3-8) AH = CP(T2-T,)
The specific heat varies with temperature and composition of the gas stream.
Therefore, Equation 3-8 is used to obtain an approximate value. For most
incineration systems, the waste gases are considered to be essentially air. For air, an
average specific heat value, Cp, is 1088 J/kg»°R (0.26 Btu/lb«°F) for typical
temperature ranges normally encountered in fume incinerators.
Equation 3-8 depicted the amount of heat required to raise a set quantity of gas
from TI to T2. Of more importance is the heat rate, q, that is required. The heat
rate is easily determined by multiplying either side of Equation 3-8 by the mass
flow rate (m) of the process gas stream. The heat rate required is given by:
(Eq. 3-9) q = rhAH = rhCp(T2 - T,)
Where: q=heat rate, watts (Btu/hr)
rh = mass flow rate of gases, kg/hr (Ib/hr)
In either form, Equation 3-9 can be used to compute the heat rate required to
raise the gas temperature from T[ to T2. Example 3-2 illustrates the use of
Equation 3-9.
These equations are simple heat balances, equating heat in to heat out. They do
not account for any heat losses in the system. Heat losses from refractory or ducting
are usually accounted for by assuming a fixed percent of the total theoretical heat
is lost. For example, if an afterburner is required to supply heat at the rate of
3 x 105 watts ( « 1 x 106 Btu/hr) and there is a 10% heat loss from the combustion
chamber, the total heat rate would have to be 3.3X 105 watts to account for the
losses. Heat is also lost to the system from the latent heat associated with vaporiza-
tion, sublimination and fusion of various components in the waste gas stream. This
is the heat which produces a change of state at a constant temperature.
The following example illustrates a method that can be used to determine the
fuel requirement for an incinerator. The example has been worked using English
units since figures and tables needed to solve the problem are also in English units.
3-13
-------
Example 3-2
The exhaust from a meat smoke house contains obnoxious odors and fumes The
company plans to incinerate the 5000 acfm exhaust stream. What quantity of
natural gas is required to raise the waste gas stream from a temperature of 90 °F to
the required temperature of 1200°F? The gross heating value of natural gas is 1059
Btu/scf and assume no heat losses.
Solution:
All calculations are based on a 1 hour time period.
First, the volume of waste gas must be corrected to standard conditions (60 °F and
1 atm).
ft3 46° + 60 6
= "84 X 105 ftVhr
tO a mass flow rate by multiplying by
(assume waste gas molecular weight is the
/ > , . . r same as air)
= (2.84X 10»ftA/lbmole\/J9_lb_\
\ hr A 379 ft3 Alb mole/
(conversion factor for an ideal gas at 60 °F)
= 2.17x10* Ib/hr
The heat rate can be determined by two methods using Equation 3-9.
1. By using the enthalpy values in Table 3-2:
H for air at 1200°F=288.5 Btu/lb
H for air at 90 °F is obtained by interpolating = 9.6- [9.6 (— )]= 7.2 Btu/lb
q = rhAH
= rfl(H-'>,2oo.-H<»>9o.)
= (2.17x10* lb/hr)(288.5 - 7.2 Btu/lb)
= 6.10X106 Btu/hr
2. An alternative method is to use an average C
For air Cp = 0.26 Btu/lb«°F
= mCAT
p
= 6.26xl06 Btu/hr
Thisvalue is not as accurate, since it was calculated using an average value
lor
3-14
-------
To compute the amount of natural gas required from the heating rate, the
available heat of the fuel (H^) must be computed using Figure 3-1.
Figure 3-1 is used by locating the flue gas temperature on the x axis (1200°F).
Read up from this point to the line for natural gas with a heating value of
1059 Btu/scf. The heat available is read from the y axis as 690 Btu/scf.
The amount of natural gas needed is:
Q-=H
riA
scf
hr 690 Btu
= 8841 S-lf
hr
Combustion Equipment Used for Control
of Gaseous Emissions
Introduction
Afterburning, incineration, or thermal oxidation are interchangeable terms used to
describe the combustion process used to control gaseous emissions. Although not
technically correct, this nomenclature is generally accepted. Afterburners and
incinerators perform different functions even though both are forms of thermal
oxidation. The term afterburner is appropriate only to describe a thermal oxidizer
used to control gases coming from a process where combustion was not complete
(Ross, 1977). An example of an afterburner in use would be an application on a
copper wire reclaiming source which incompletely burns off the rubber coating.
Incineration or thermal oxidation are not necessarily afterburning. Each is a
primary burning process that is installed to control effluent pollutants which are
combustible. Incinerators are used to combust solid, liquid and gaseous materials.
When used in this manual, the term incinerator will refer to controlling gaseous
emissions of organic vapors. Examples of sources using incineration are curing and
drying ovens, chemical processes, petroleum refining processes, and food processing
plants. In this chapter the terms incinerator and thermal oxidizer will be used
interchangably.
Equipment used to control waste gases by combustion can be divided into three
categories: direct combustion or flaring, thermal oxidation, and catalytic oxida-
tion. A direct combustor or flare is a device in which air and all the combustible
waste gases react at the burner. Complete combustion must occur instantaneously
since there is no residence chamber. Therefore, the flame temperature is the most
important variable in flaring waste gases. In contrast, in thermal oxidation, the
combustible waste gases pass over or around a burner flame into a residence
chamber where oxidation of the waste gases is completed. Catalytic oxidation is
3-15
-------
very similar to thermal oxidation. The main difference is that after passing through
the flame area, the gases pass over a catalyst bed which promotes oxidation at a
lower temperature than does thermal oxidation.
Direct Combustion or Flaring
Direct combustion or flaring is used for the disposal of intermittent or emergency
emissions of combustible gases from industrial sources. Safety and health hazards at
or near the plant can be eliminated by using flares to prevent the direct venting of
these emissions. Flares have been used mainly at oil refineries and chemical plants
which handle large volumes of combustible gases.
Flares are simply burners that have been designed to handle varying rates of fuel
while burning smokelessly. In general, flares can be classified as either elevated or
ground level. The reason for elevating a flare is to eliminate any potential fire
hazard at ground level. Ground level flares must be completely enclosed to conceal
the flame. Either type of flare must be capable of operating over a wide range of
flow rates in order to handle all plant emergencies. The range of waste gas flows
within which a flare can operate and still burn efficiently is referred to as the turn-
down ratio. Flares are expected to handle turndown ratios of 1000:1, while most
industrial boilers seldom handle more than a 10:1 turndown ratio (Gottschlich,
1977). For example, a flare should be capable of maintaining complete combustion
for waste gas flow rates ranging from 20,000 mVh to 20 mVh.
Although the flare is designed to eliminate waste gas stream disposal problems, it
can present safety and operational problems of its own. Some of the problems
associated with operation of a flare system are:
1. Thermal radiation: Heat given off to the surrounding area may be
unacceptable.
2. Light: Luminescence from the flame may be a nuisance if the plant is located
in an urban area.
3. Noise: Mixing at the flare tip is done by jet Venturis which can cause excess
noise levels in nearby neighborhoods.
4. Smoke: Incomplete combustion can result in toxic or obnoxious emissions.
5. Energy consumption: Flares waste energy in two ways. First by keeping the
pilot flame constantly lit and secondly, by the potential recovery value of the
waste gas being flared.
For flaring to be economically feasible, the waste gas usually must supply at least
50% of the fuel value to combust the mixture. When the heat content of the waste
gas is below 4.28 x 10s kj/m3 (115 Btu/ft3), injecting an additional gas with a
higher heating value is necessary to achieve complete combustion. This type of
system is referred to as a fired or endothermic flare. The cost of the additional fuel
for endothermic flaring can be considerable. To conserve energy, some companies
have used other waste gases in place of conventional fuels.
3-16
-------
Elevated Flares
A typical elevated flare is composed of a system which first collects the waste gases
and passes them through a knockout drum to remove any liquids. Water seals or
other safety devices are placed between the knockout drum and the flare stack to
prevent a flashback of flames into the collection system. The flare stack is essen-
tially a hollow pipe that may extend to a height of 100 meters or more. The
diameter of the flare stack determines the volume of waste gases that can be
handled. At the top of the stack is the flare tip which is comprised of the burners
and a system to mix the air and fuel.
Smokeless combustion is accomplished by proper design of the flare tip.
Smokeless flares require a system allowing for intimate mixing of waste gases and
air. A number of flare tip designs provide good mixing characteristics over a wide
range of waste gas flow values and still have excellent flame holding capabilities.
Steam jets have proved to be one of the most effective ways to mix air and waste
gases. An example of a flare tip designed for steam injection is illustrated in Figure
3-4. In addition to increasing turbulence, steam injection reduces the partial
pressure of the waste gas, thus reducing polymerization reactions. The steam also
reacts with the gases to produce oxygenated compounds which readily burn at
lower temperatures. Other devices used to reduce smoke are high pressure fuel
gases, water sprays, high velocity swirl fired burners, and electric air blowers
(Straitz, 1980). In addition to these devices, shields are also used on elevated flares
to protect the flame zone from atmospheric conditions. This shield also helps
reduce noise and visibility problems associated with flares.
Steam injection point
Waste gas
retention ring
Steam distribution
ring
Flare stack
Pilot assembly
Figure 3-4. Smokeless flare tip.
3-17
-------
As previously mentioned, steam reacts with intermediate combustion products to
form compounds that burn readily at lower temperatures. One of the main reac-
tions taking place is a hydrolysis reaction referred to as the water-gas reaction This
reaction produces hydrogen gas which helps flare operation. The simplified water-
gas shift can be written as:
Steam-to-hydrocarbon mass ratios are usually determined by the molecular weight
and concentration of the unsaturated organic compounds (alkenes and alkynes) in
the waste gas. Steam requirements generally range between 0.05 and 0 30 kg of
steam per kg of waste gas (Gottschlich, 1977). The steam is automatically injected
proportionally to the flow rate of the waste gases.
Enclosed Ground Flares
Manufacturers design a number of different ground flares. Most ground flares con-
sist of multiple burners enclosed within a refractory shell. The shell encloses the
name to elimmate noise, luminescence, and safety hazards. The waste gas is
introduced through a jet or venturi to provide turbulent mixing. The term ground
flare refers to locating the flare tip at ground level. The flare system still requires a
stack for proper release of the effluent gases. Figure 3-5 shows a ground flare
installed at Nippon Steel Company in Oita City, Japan (Straitz, 1980). The flare is
composed of two chambers designed to combust eight different gases and a liquid-
waste stream. 4
Burners
Accoustical insulation
Liquid-waste
atomizing injectors
Source: Straitz, 1980.
Figure 3-5. Ground flare.
3-18
-------
Ground level flares normally have a much higher capital cost than elevated
flares. The cost of a ground flare depends on the size of the enclosure. The size of
a ground flare is directly proportional to the volume of vapors it must handle.
Ground flares can be designed for efficient combustion without steam injection.
Eliminating the use of steam injection can significantly lower operating costs as
compared to elevated systems (Straitz, 1980). Some plants have incorporated both
designs. A ground flare is used for normal or intermittent operation and a large
elevated flare is only used to control emergency releases of large quantities of gases.
Incinerators
Thermal oxidizers or organic vapor incinerators refer to any device that uses a
flame (temperature) combined with a chamber (time and turbulence) to convert
combustible material to carbon dioxide and water. An incinerator usually consists
of a refractory-lined chamber that is equipped with one or more sets of burners. A
typical incinerator is depicted in Figure 3-6. The contaminant-laden stream is
passed through the burners where it is heated above its ignition temperature. The
hot gases then pass through one or more residence chambers where they are held
for a certain length of time to ensure complete combustion. Depending on the par-
ticular needs of the system, additional fuel and/or excess air can be added through
the burners. Also, since the flue gases are discharged at elevated temperatures, a
system to recover the heat may be included. With the rising costs of fuel and
solvents, heat recovery devices are becoming an integral part of many incineration
systems.
Plenum
Figure 3-6. Typical thermal incinerator (UOP raw gas burner).
Incinerators on industrial processes are most often used to control gas streams
with a low concentration of organic vapors. The concentration of vapors delivered
to an incinerator through process ductwork is limited by the LEL of the mixture.
Safety codes usually limit combustible vapor concentrations in ducts to a maximum
of 25% of the LEL. By using process control monitors and alarm systems, some
plants have been allowed to go up to 50% of the LEL. Supplemental fuel is
required to ignite the waste gas stream and bring it up to the proper operating
3-19
-------
temperature. The amount of fuel is then decreased proportionally to the heating
value supplied by the waste gas stream. Ideally, after initial startup, little or no
auxiliary fuel would be needed.
Incinerators operate at temperatures between 700 and 800 °C (1300 and 1500°F)
with a residence time of 0.1 to 0.5 seconds. The residence time is determined by
the size of the combustion chamber and is measured after the required
temperature has been reached. Typical operating temperatures of thermal oxidizers
for various processes are listed in Table 3-3 (EPA, 1972). As noted previously
temperature and residence must be discussed together since changing one affects
the other. By raising the temperature, the residence time for complete combustion
is reduced and vice versa. However, temperature is the more important process
variable. Figure 3-7 illustrates the effect of both temperature and residence time on
the percent destruction of pollutants (oxidation rate). Depending on the initial
temperature, small increases in temperature can bring dramatic increases in pollu-
tant destruction. For example, for a 0.01 second residence time increasing the
temperature from 1200 to 1400°F, the percent destruction is doubled from approx-
imately 50 to 100%. At 1200°F, the residence time must be increased ten times
(from 0.01 second to over 0.1 second) for the same increase in percent pollutant
destruction.
Table 3-3. Typical waste gas incinerators' operating temperatures (°F).
Industry
Asphalt blowing
Biological control, fermentation
Carpet laminating
Coffee roasting
Coil coating, sheet coating, metal decorating
Core ovens, foundry
Coating, engraving
Cloth carbonization
Deep fat fryers
Gum label drying oven
Mineral wool, fiberglass curing
Odor control (general) sludge off-gas
Hardboard tempering
Oil and grease smoke (metal chip recovery, heat
quench baths, tempering)
Paint bake ovens
Paper manufacture — sulfite digester off-gas
Pipe wrapping
Rendering plants
Rubber products
Petroleum refining and products
Printing, lithographing
Smelting, refining, metal recovery, wire burnoff
Smokehouse operation
Solvent control
Varnish cookers, resin kettles
Vinyl plastisol curing
Wood milling
Wire enameling
Phthalic anhdyride
Textile drying oven
LAAPCD**
recommendation
1200-1400
1400
1800
1200
1300-1500
1200-1400
1200-1500
1400
1200
1200
1300-1500
1200
1200-1400
-
Survey data
1000-2000
110-1250
1200-1500
1200-1500
2000
1000*-1450
1000*
1100-1250
1200
900M600
1100-1500
1200-1300
1300-1400
1300-2000
1300-1500
1300-1650
800*-1200
1000*-1500
1200
1250-1440
1350
Literature
-
1050
1300
-
~
1 OKrt
IzoO
1310
1300-1425
1200
1240
1350
1200
1300
1200-1400
900*
1 9flfi
IZUU
1300-1350, 1400
temperature generally for odor, smoke control, not true organic vapor destruction.
*LAAPCD is the Los Angeles Air Pollution Control District.
3-20
-------
o
C
ra
o
O,
100
80
60
40
20
I
I
600
800
1000
1200 1400
Temperature, °F
1600 1800
Source: Rolke, 1972.
Figure 3-7. Effects of temperature and residence time
on rate of pollutant oxidation.
Besides temperature and residence time, the concentration of the pollutant in
the waste gas stream also affects operation of the incinerator. First, the concentra-
tion of the contaminant dictates the amount of supplemental fuel required.
Initially, supplemental fuel must be added to start the oxidation reaction. As the
temperature rises, the rate of oxidation of the contaminant increases until the reac-
tion becomes self-sustaining. The amount of fuel can then be decreased. A certain
amount of fuel is normally burned to ensure stable operation. Secondly, there is a
problem of flame quenching. To avoid noncombustible mixtures, the entire
amount of waste gas and fuel cannot usually be mixed at the burner. It would
require an inordinate amount of fuel to bring the entire waste gas stream within
combustible limits. Only part (about one-half) of the waste gas stream is mixed
with the fuel at the burner. The remainder of the waste gas stream must be mixed
with the hot products of combustion downstream of the flame to avoid quenching
the flame. Unless adequate mixing of the waste gas stream and hot products of
combustion is accomplished, incomplete combustion products will be exhausted.
3-21
-------
Types of Burner Arrangements
The design of an incinerator must allow adequate mixing of the fuel and waste gas
streams. A number of different burner designs provide the required mixing without
quenching the flame. Burners are classified as either discrete or distributed (EPA,
1973). The main distinction between the two types of burners is the number and
size of flames in each type. The discrete burner produces a single large flame
plume into which the waste gas stream must be blended. The distributed burner
system produces an array of smaller flames. The waste gas stream is divided and
then flows around these flames. In distributed burners more of the waste gas
stream is passed over or around the flame. This allows distributed burners to com-
plete mixing of the waste gases and hot products of combustion in shorter distances
than in discrete burners. Distributed burners also utilize the oxygen in the waste
gas stream more efficiently, reducing the amount of oxygen supplied through the
burner. This in turn reduces the need for additional fuel to heat the air/waste gas
mixture to the required oxidation temperature.
Three types of distributed burner systems are the line burner, the multijet, and
the grid burner. All three of these systems are comprised of a manifold system, to
inject the fuel, and a mixing or profile plate which regulates air flow to the flame
area. Fuel (usually natural gas) enters the burner area of the line burner (Figure
3-8) through holes in manifold pipes. The pipes are placed across the duct at 0.3
meters (1.0 foot) intervals (EPA, 1972). Waste gas enters the burner area through
holes in the profile plate. The plate allows only a portion of the waste gas, usually
50%, to directly contact the flame. The remaining portion of the gas bypasses the
flame area and is mixed with the hot combustion gases downstream of the flame.
Figure 3-9 illustrates an example of one type of mixing plate. The plate is attached
to the manifold pipes and forms a "v" shaped trough; part of the waste gas enters
the flame area through the holes, while the remaining portion of waste gas passes
around the plate.
3-22
-------
Fuel
Combustion
chamber
Waste gas and air
Figure 3-8. Distributed (line) burner.
Burners
Burners
Fuel
Figure 3-9. Mixing plate for waste gas and hot combustion products.
3-23
-------
Multijet burners are very similar to line burners. Multijet burners have a single
fuel inlet which is manifolded to produce numerous individual flames. Figure 3-10
shows the Hirt multijet burner system. In this system, waste gases flowing to the
flame area are controlled by a profile or mixing plate. The profile plate allows a
portion of the waste gas stream to flow behind it and mix with the fuel to supply
the combustion area. The remaining gases must pass around in front of the profile
plate. Air flow to the burner flame area is adjusted by moving the profile plate
either closer or further away from the burners. Adjusting the profile plate on the
multijet system is relatively simple and can be accomplished while the unit is still in
service. Adjusting the mixing plate on the line burner is not as simple, because the
unit must be taken off line. The line burner, however, has the advantage of being
much smaller. The line burner accomplishes mixing over a relatively short
distance, a few inches, as compared to the multijet which requires several feet.
Adjustable gap
Mixing plate
Multijet burners
Fuel
Waste gas
Figure 3-10. Multijet burner.
The grid burner also uses a manifold system to introduce the fuel. Mixing is
accomplished by passing the waste gas stream through a slotted grid as shown in
Figure 3-11. The slots are approximately 2.5 by 7.5 cm (1 by 3 in.) (EPA, 1972).
The grid allows for extremely fast mixing of the fuel and waste gas. There is,
however, very little control to allow for adjusting flow rates since the grid is sta-
tionary. Grid burners are best suited for applications where waste gas flow ratios
are relatively constant.
3-24
-------
Fuel
Flame area
Waste gas
Figure 3-11. North American flame grid burner.
Problems can arise using any of the three types of distributed systems. Some of
these are:
1. Natural gas is the only fuel source that can be used. If solid or liquid par-
ticles are present or formed during combustion, they can plug the tiny holes
in the fuel manifold system. This is known as burner fouling.
2. The oxygen content of the waste gas stream must be at least 16%. Oxygen for
combustion is supplied by the waste gas stream. If the oxygen content is not
16% or greater, incomplete combustion can occur.
3. Mixing, profile, and grid plates must be able to withstand high temperatures.
This adds to the capital equipment costs.
3-25
-------
To avoid some of the problems associated with distributed burners, especially
fouling, discrete burners are sometimes used. The discrete burner is simply one
large burner (Figure 3-12). Many other burner arrangements are classified as
discrete burners. They can range from a simple gas ring, comparable to those used
on home stoves, to a torch burner (Figure 3-6). In this illustration, mixing the pro-
cess waste gas and hot products of combustion is accomplished by using a conical
mixing plate.
Waste gas
Fuel and air
Conical mixing plate
Figure 3-12. Discrete burner.
In most discrete burners, mixing must usually be provided downstream of the
flame. Baffles are most commonly used to provide the required turbulence.
However, baffles increase the pressure drop across the incinerator. Unless installed
correctly, baffles can cause dead zones which actually decrease the degree of
mixing. Two main types used in incinerators are the bridge wall baffle and the
ring and disc baffle illustrated in Figure 3-13 and Figure 3-14. The bridge wall
baffle is a wall placed across the combustion chamber. The bridge wall technique
is most effective when pairs of baffles are used. The ring and disc baffles consist of
a series of rings inserted in the middle along with discs on the side walls of the
chamber.
Figure 3-13. Bridge wall baffle.
3-26
-------
Figure 3-14. Ring and disc baffle.
Catalytic Oxidation
A catalyst is a substance which causes or speeds a chemical reaction without itself
undergoing a change. In catalytic incineration, a waste gas is passed through a
layer of catalyst known as the catalyst bed. The catalyst causes the oxidation reac-
tion to proceed at a faster rate and lower temperature than is capable in thermal
oxidation. Catalytic incinerators operating in a 370 to 480°C (700 to 900°F) range
can achieve the same efficiency as a thermal incinerator operating between 700 and
820°C (1300 and 1500°F). This can result in a 40 to 60% fuel savings.
Catalytic reactions can be classified as either homogeneous or heterogeneous.
Homogeneous reactions occur throughout the bulk of the catalyst, while
heterogeneous reactions occur only on the surface of the catalytic material. In air
pollution control applications, all reactions are heterogeneous. The oxidation reac-
tion of the organic vapors occurs only on the surface of the catalyst. It should be
noted that catalytic reactions produce the same end products (CO2 and H2O) and
liberate the same heat of combustion as does thermal incineration.
A heterogeneous catalyst reaction proceeds through a series of five basic steps:
1. Organic compounds in the waste gas must first diffuse from the bulk of the
vapor to the surface of the catalyst.
2. Organic compounds then adsorb onto the surface of the catalyst.
3. Organic compounds then react (oxidize).
4. New compounds then desorb after reacting.
5. New compounds then diffuse and mix back into the bulk of the exhaust air
stream.
The most effective and commonly used catalysts for oxidation reactions come
from the noble metals group. Platinum either alone or in combination with other
noble metals is by far the most commonly used. Desirable characteristics of
platinum are that it gives a high oxidation activity at low temperatures, is stable at
high temperatures, and is chemically inert. Palladium is another noble metal which
exhibits these properties and is sometimes used in catalytic incinerators.
3-27
-------
Since catalytic oxidation is a surface reaction, the noble metal is coated onto the
surface of a cheaper support material. The support material can be made of a
ceramic or a metal such as alumina, silica-alumina or nickel-chromium. The sup-
port material is arranged in a matrix shape to provide: high geometric surface
area; low pressure drop; uniform flow of the waste gas through the catalyst bed;
and a structurally stable surface (EPA, 1972). Structures which provide these
characteristics are pellets, a honeycomb matrix, or a mesh matrix. Figure 3-15
shows a typical honeycomb catalyst module which is the most common. The sup-
port material for these is usually ceramic, but can be metal.
Figure 3-15. Typical honeycomb catalysts (metallic or ceramic).
A schematic of a catalytic incinerator is shown in Figure 3-16. Catalytic
incinerators consist of a preheat section (burner area), where part of the waste gas
is raised to operating temperature. The burners are the same as those used for
thermal incineration, with the majority being distributed burners. The remaining
portion of the waste gas is mixed with the hot products of combustion before
passing over the catalyst bed. This ensures a homogenous waste gas and
temperature mixture as it passes over the bed. After passing over the bed, the hot
flue gases may be sent to a heat recovery system.
3-28
-------
Heat exchanger tubes
Catalyst
Figure 3-16. Catalytic incinerator.
Catalyst bed depth controls the pressure drop across the incinerator. Typically,
the volume of catalyst required for 85 to 95% conversion of all organic compounds
in 0.01 to 0.05 seconds is between 0.03 to 0.14 m3 of catalyst per 1000 m/s of
waste gas (0.5 to 2.0 ft3 of catalyst per 1000 cfm of waste gas). As a general rule,
the higher the molecular weight of the compound, the more readily it is oxidized.
Normal pressure drops through catalytic incinerators are between 62 and 125 Pa
(0.25 and 0.5 in. H2O). If the exhaust does not have sufficient draft to overcome
this pressure drop, a blower may have to be installed.
To ensure proper operation, the inlet and outlet temperatures to the catalyst sec-
tion are monitored. This verifies that the temperature is sufficient to achieve the
required conversion and also ensures protection of the catalyst from excessive
temperatures. Since catalytic incinerators operate at lower temperatures than ther-
mal units, less refractory brick and small chamber volumes can be used. This
reduction in size of the device reduces installation costs.
Operating Limitations of Catalytic Incinerators
Catalytic incinerators usually cannot be used effectively on waste gas streams which
contain particulate matter. Paniculate matter that deposits on the surface of the
catalyst prohibits the organic compounds from being adsorbed. Coating of the
catalyst surface in this manner is referred to as fouling of the catalyst. Oil droplets
can also foul the catalyst bed unless they are vaporized in the preheat section. By
periodically cleaning and washing the catalyst, 90% of its activity can be restored
(EPA, 1972). Maintenance of this type, however, adds significantly to the cost of
operating the unit.
Certain metals can chemically combine with or alloy to the catalyst, thereby
making it useless. Deactivation in this manner is called catalyst poisoning. Catalyst
poisons can be divided into two categories: fast acting poisons which include
3-29
-------
phosphorus, btsmuth, arsenic, antimony, and mercury; and slow acting poisons
which include zmc, lead and tin. Catalysts are more tolerant of the slow acting
poisons, particularly at temperatures below 540 °C (1000°F). At sufficiently hieh
temperatures (> 590°C), even copper and iron are capable of alloying to platinum
reducing its activity.
Sulfur and halogen compounds act as poisons, but not to the extent of the
metals. They are essentially reaction inhibitors called suppressants. Their chemical
interaction with the catalyst is reversible. Once the halogen or sulfur compound is
removed, catalyst activity is restored to normal.
All catalysts deteriorate with normal use. Gradual loss of the catalyst material
can occur from erosion, attrition, and vaporization. High temperatures can also
accelerate catalyst deactivation. Loss of catalyst activity due to high temperature is
known as thermal aging, and causes very rapid catalyst deterioration. Even short
term temperatures above 820°C (1500°F) can cause a near total loss of catalyst
activity. With proper monitoring of operating temperatures, a catalyst bed can be
expected to last from three to five years before it must be replaced. The percent
destruction of pollutants decreases with increasing thermal aging. Adjustment to
operating conditions (temperature, residence time, etc.) may need to be made to
ensure that the exhaust continues to meet emission limits.
Comparison of Thermal vs. Catalytic Incinerators
The major difference between thermal and catalytic incinerators is that complete
combustion can be achieved at much lower temperatures in a catalytic incinerator
The reduced temperatures cut the cost of fuel usage by 40 to 60%. Operation at
lower temperatures also decreases the construction costs. Lighter materials of con-
struction can be used in catalytic units. This also makes installation easier and less
costly. In terms of overall equipment costs, for small capacity units, up to 4 7 or
5.7 rnVs (10,000 or 12,000 dm), the purchase costs for either unit is essentially
equal (EPA, 1972). For larger units that must be custom designed, thermal
incinerators are usually less expensive, but this depends on the type of heat
recovery system.
The main problem in catalytic incineration is the reduction or loss of catalyst
activity. Loss of catalyst activity occurs due to fouling by paniculate matter or sup-
pression or poisoning due to other contaminants in the waste gas stream. In order
to effectively use catalytic incineration, these contaminants must be removed from
the waste gas stream. Removing these contaminants would require additional
equipment which adds greatly to the cost of the system. Finally, all catalysts must
periodically be replaced due to thermal aging.
Process Boilers Used as Incinerators
An alternative to installing a thermal or catalytic incinerator would be to combust
the waste gases in an existing plant or process boiler. This would avoid the capital
cost of new equipment and may help to reduce present fuel consumption. Process
and plant boilers are normally designed to operate in excess of 980 °C (1800°F)
3-30
-------
with a flue gas residence time of 0.5 to 3.0 seconds. These conditions exceed those
recommended for thermal incinerators. However, a number of additional condi-
tions must be satisfied before waste gases can be properly disposed of in this
manner.
The following criteria must be considered before a process boiler can be used as
an incinerator (EPA, 1977):
1. The waste gases must be almost completely combustible. If solid particles
are present in the waste gases, or formed by incomplete combustion, they can
foul heat exchanger surfaces in the boiler, thus reducing boiler efficiency.
Particulate matter may also cause boiler emissions to exceed applicable emis-
sion regulations. The costs of increased maintenance of the boiler and/or con-
trol of particulate matter may well exceed the purchase price of an
incinerator.
2. The waste gas should, preferably, constitute only a small fraction of the air
requirements of the boiler. If the volume of the waste gas is large, special
attention must be paid to the oxygen balance, mixing, and continuation of
the air flow in the boiler when the process is shut down.
3. The oxygen concentration of the contaminated gas stream should be close to
that of air to avoid incomplete combustion. Incomplete combustion produces
tars that coat heat exchanger surfaces, reducing boiler efficiency.
4. The boiler must operate at all times when incineration is required.
5. The waste gas must be free of compounds, such as halogenated hydrocarbons,
that accelerate corrosion of the boiler.
6. Baffling may be required in the combustion chamber to ensure adequate
mixing and combustion of the waste gas.
7. If the boiler-firing rate varies greatly, it may be necessary to install a small
auxiliary boiler that will operate under steady load conditions.
To date, not many industries have been successful in using plant or process
boilers as incinerators. Petroleum refineries, which have numerous waste gas
streams and process boilers, are one of the few industries which have been able to
incinerate waste gases in process boilers.
Heat Recovery Systems
Due to the rising cost and limited availability of fuels, heat recovery systems are
becoming an integral part of most incinerators. Heat recovered from hot flue gas
can be used in two ways to reduce energy consumption. One way is to preheat the
cooler waste gas from the process entering the incinerator, referred to as primary
heat recovery. The second way is to use the recovered heat in another process in
the plant, such as a drying oven, referred to as secondary heat recovery. A number
of systems are designed to recover heat in either one or in some cases both
manners. To date, however, no single heat recovery system works universally well
on all processes.
3-31
-------
Specific terms are used in this chapter to describe heat recovery. Waste gas will
always refer to the contaminant gas stream entering the incinerator. Flue gas refers
to the hot products of combustion that are leaving the incinerator. Also, note that
the terms hot and cold are relative, describing the direction that heat is being
transferred. For example, hot flue gas may be used to heat the cold waste gas
stream (even though the waste gas may be at 150°C).
Heat recovery devices are classified by their heat transfer method. Heat can be
transferred either directly or indirectly. Mixing all or part of the hot flue gases with
the waste gas is an example of direct heat recovery. Indirect heat recovery devices
use heat absorbing or heat transferring materials to recover heat. Indirect heat
recovery devices are either recuperative or regenerative. Recuperative heat recovery
devices use heat transferring material, such as a tube or bank of tubes with hot gas
on one side and cold gas on the other. A home hot water heater is an example of
this system. Regenerative heat recovery devices use heat absorbing materials, such
as stones or ceramic packing, to store heat from a hot gas stream then give it up to
a cooler stream.
To be effective, heat recovery devices must take as much heat as possible from
the hot flue gas and transfer it to the cooler waste gas stream. There are two
temperature limiting factors to be aware of when trying to maximize this heat
transfer. First, the degree of heat recovery is limited by the dew point of any gases
in the hot flue gas stream. If the hot flue gas is cooled to below the dew point of
any constituent (gas), corrosive acids can be formed. These acids can cause
excessive damage to any heat recovery device or associated ductwork. One solution
is to use construction materials resistant to acid corrosion. However, these are
expensive and add greatly to system cost. The practical solution is to avoid excess
cooling of the hot flue gas (below the dew point). For example, the dew point of
sulfur-free fuels (i.e. natural gas) is approximately 60°C (140°F). Therefore, this
temperature sets the lower limit of heat that is possible to be recovered.
Secondly, heat recovery can also be limited by the autoignition temperature of
the waste gas stream. If the waste gas is preheated to its autoignition temperature
an explosion or fire could occur. The autoignition temperatures of most industrial
hydrocarbons fall between 370 and 540°C (700 and 1000°F). Heat recovery systems
are capable of preheating waste gas streams above these temperatures. Unless other
uses for the heat are found, safety restrictions will not permit utilizing all of the
available heat energy.
Recuperative Heat Recovery Devices
Recuperative heat recovery devices are simply tubular heat exchangers (Figure
3-17). In these devices the cold waste gas usually passes on the inside of the tube,
while the hot flue gas passes over the outside of the tubes. The effectiveness of
these devices is limited by the thermal resistance to heat transfer through the tubes.
Using circular tubes maximizes the amount of surface area available for heat
transfer. These are currently the most frequently used heat recovery devices since
they contain no moving parts.
3-32
-------
Preheated waste gas
To stack
exhaust
Figure 3-17. Recuperative heat exchanger (shell-and-tube).
In comparing different heat recovery devices, it is mandatory to have a common
basis for judging performance. One method used to judge the performance of heat
recovery systems is by the heat transfer effectiveness (E) measured by (EPA, 1972):
(Eq. 3-10)
E=-a_ =
gas en
— T
ttring A//u*
gas leaving}
Heat transfer effectiveness is measured as the ratio of the amount of heat
actually recovered (q) to the maximum amount of heat (qma*) that potentially could
be recovered. Examining the right side of Equation 3-10, it can normally be
assumed that the mass flow (m) and specific heat (Cp) of both the waste gas and
flue streams are equivalent. For this case, the heat recovered depends only on the
temperature difference of the streams entering and leaving the device.
For single pass, tubular heat exchangers as shown in Figure 3-17, the effec-
tiveness is usually between 40 and 50% (EPA, 1972). By arranging multiple passes
of either hot or cold streams, heat exchanger effectiveness can be increased. By
passing the cold waste gas over the tubes twice, the effectiveness can be increased
to approximately 60%; while for a three pass system, the effectiveness can be raised
3-33
-------
to 70% (Mueller, 1977). Figure 3-18 shows a typical double pass heat exchanger
system. Although the effectiveness increases by adding extra passes or stages, it also
adds to the total pressure drop. The pressure drop across a typical tubular heat
exchanger is between 1 to 2 kPa per stage (4 to 8 inches of H2O per stage).
Figure 3-18. Double pass recuperative heat exchanger.
The main disadvantage of the recuperative heat recovery devices are that they
are easily fouled by heavy smoke or sticky paniculate matter. Fouling occurs when
material coats the tubes, reducing the rate of heat transfer. If fouling is a problem,
frequent cleaning of the tubes may be required. When contaminants are present,
the contaminated stream is usually passed through the inside of the tubes. Cleaning
the inside of a tube is much easier than cleaning the outside. Materials of construc-
tion for heat exchangers must be able to withstand constant operation at high
temperatures. Heavy duty, high quality materials of construction must be used or
thermal expansion can cause warpage and thus cause welds to separate. If this occurs,
contaminants can bypass the incinerator and be exhausted directly to the
atmosphere. In addition, the waste gas must not be preheated above its auto-
ignition temperature or cooled below its dew point to avoid damage to the device.
Regenerative Heat Recovery Devices
Regenerative devices use absorbing material that stores heat from hot flue gas for
later use. These devices operate on a cycle, part of the time the absorbing material
is being heated, then that portion of the bed is moved to regeneration or heat
recovery cycle. Regenerative devices can recover a greater percent of heat than
recuperative devices. Regenerative devices have only one gas-solid heat resistance
3-34
-------
boundary (between hot gas and absorbing material) to pass through. In
recuperative devices heat must pass through two transfer boundaries (the inside and
outside of the tube). The two most prevalent regenerative devices are the rotary
exchanger and the packed bed.
The rotary or wheel heat exchanger consists of a wheel with a large surface area
(Figure 3-19). Hot flue gas flows through one portion of the wheel while the cool
waste gas flows in the remaining portion. The surface of the wheel is constructed of
corrugated sheet metal plates. The plates are arranged in a honeycomb matrix
form to give both maximum heat transfer and air flow between the plates.
Rotating the wheel allows one of its sides to heat up, then this heat is transferred to
the cool side. Rotary wheels of this type are capable of achieving 80% effectiveness
and have very low pressure drops (Bethea, 1978).
Hot flue gas
Preheated
waste gas
To stack
exhaust
Process
waste gas
Figure 3-19. Regenerative heat wheel.
Disadvantages associated with use of the heat wheel are clogging and thermal
wear. The small passageways between the plate of the wheel can become plugged
by heavy smoke or sticky particulate matter in the waste gas stream. Isolating seals
control air leakage between the hot and cold sections. These seals are susceptible to
wear and warpage due to high temperatures and constant movement. Seal failure
3-35
-------
allows contaminants to bypass the incinerator. Leakage from these seals can be
minimized in some instances by arranging blowers to provide a higher pressure on
the flue gas side of the wheel so that any leaks will be directed into the waste gas
stream.
The other regenerative heat recovery device is a packed bed. One design consists
of a number of towers packed with a heat absorbing material. The towers surround
a central combustion chamber. A typical regenerative heat recovery device
manufactured by REECO, Inc. is shown in Figure 3-20. Initially the waste gas
passes into the combustion chamber where it is incinerated. Hot flue gas passes
from the combustion chamber through a packed bed, heating the bed. Once a
packed bed is heated, flow switching valves direct the cold waste gas through the
hot bed which preheats it to combustion temperatures. The cycle continues with
the hot flue gas heating another packed section. At steady state conditions it is
possible to operate the REECO heat recovery device without any auxilary fuel. This
capability depends on the concentration and heating value of the combustible com-
pounds in the waste gas.
Packed
bed
Source: REECO, Inc.
Figure 3-20. Regenerative heat recovery device.
3-36
-------
The REECO device has been marketed only for the past few years. Although
there is not a lot of data on these systems, the installed units seem to be operating
well. These are large units, approximately 10 meters high by 7 meters in diameter.
Therefore, they are economical only for large industrial applications. They are
designed for high effectiveness, between 70 and 90%. Due to their unique design,
they are self-cleaning and they also eliminate the hazard of autoignition in
preheating the waste gas. '
Economic Considerations of Heat Recovery
A recent EPA study of ten heat recovery systems noted that the payback time for
these systems ranged from a low of 0.8 to a high of 10.9 years (EPA, 1978). Most
of the data was grouped between one and two years. With the increasing cost of
fuel, the financial benefit of heat recovery devices will increase until it is
uneconomical to operate without them.
References
Bethea, R. M. 1978. Air Pollution Control Technology. New York: Van Nostrand
Reinhold.
Environmental Protection Agency (EPA). 1972. Afterburner Systems Study.
EPA-R2-062. Research Triangle Park, NC.
Environmental Protection Agency (EPA). 1973. Air Pollution Engineering Manual,
AP40. Research Triangle Park, NC.
Environmental Protection Agency (EPA). 1977. Controlling Pollution from the
Manufacturing and Coating of Metal Products. Research Triangle Park, NC.
Environmental Protection Agency (EPA). 1978. Study of Systems for Heat
Recovery from Afterburners. Industrial Gas Cleaning Institute, Contract
No. 68-02-1473 (TASK 23). Research Triangle Park, NC.
Gottschlich, C. F., 1977. Combustion. In Air Pollution Vol. IV Engineering
Control of Air Pollution, A. C. Stern, ed. New York: Academic Press.
Mueller, J. H. 1977. Heat Recovery. In Air Pollution Control and Design Hand-
book, Part 1. P. N. Cheremisinoff and R. A. Young, eds. New York: Marcel
Dekker, Inc.
North American Combustion Handbook. 1965. North American Manufacturing
Co. Cleveland, Ohio.
Ross, R. D. 1977. Thermal Incineration. In Air Pollution Control and Design
Handbook, Part 1. P. N. Cheremisinoff and R. A. Young, eds. New York:
Marcel Dekker, Inc.
3-37
-------
Snape, T. H. 1977. Catalytic Incineration. In Air Pollution Control and Design
Handbook, Part 1. P. N. Cheremisinoff and R. A. Young, eds. New York:
Marcel Dekker, Inc.
Straitz, J. F. 1980. Flaring with Maximum Energy Conservation. Pollution
Engineering. 12:47-49.
Waid, D. E. 1974. How to Design Air Pollution Equipment that Will Not Become
Obsolete. Combustion. 45:4-12.
3-38
-------
Chapter 4
Absorption
Introduction
Absorption refers to the transfer of a gaseous component from the gas phase to a
liquid phase. More specifically, in air pollution control, absorption involves the
removal of objectionable gaseous contaminants from a process stream by dissolving
them in a liquid.
Some common terms used when discussing the absorption process follow:
• Absorbent: the liquid, usually water, into which the contaminant is absorbed.
• Absorbate or solute: the gaseous contaminant being absorbed, such as SO2,
H2S, etc.
• Carrier gas: the inert portion of the gas stream, usually air, from which the
contaminant is to be removed.
• Interface: the area where the gas phase and the absorbent contact each other.
• Solubility: the capability of a gas to be dissolved in a liquid.
Absorption is a mass transfer operation. Mass transfer can be compared to heat
transfer in that both occur because a system is trying to reach equilibrium con-
ditions. For example, in heat transfer, if a hot slab of metal is placed on top of a
cold slab, heat will be transferred from the hot slab to the cold slab until both are
at the same temperature (equilibrium). In heat transfer, the process continues as
long as a temperature differential or gradient exists. In absorption mass instead of
heat is transferred, and instead of occurring due to a temperature change, absorp-
tion occurs as a result of a concentration difference. Absorption continues as long
as a concentration differential exists between the absorbent and the gas from which
the contaminant is being removed. In absorption, equilibrium is not as easily
defined as in heat transfer, since the concentration difference depends on the
solubility of the solute.
Absorption devices used to remove gaseous contaminants are referred to as
absorbers or wet scrubbers. Wet scrubbers are also used to remove particulate
matter from gas streams. Wet scrubbers usually cannot be operated to optimize
simultaneous removal of both gases and particulate matter. In designing absorbers
for gaseous emissions, optimum mass transfer can be accomplished by:
• providing a large interfacial contact area.
• providing good mixing between gas and liquid phases.
• allowing sufficient residence or contact time between the phases.
• ensuring a high degree of solubility of the contaminant in the absorbent.
4-1
-------
Mechanism of Absorption
To remove a gaseous contaminant by absorption, the contaminant-laden exhaust
stream must be passed through (contacted with) a liquid. Figure 4-1 illustrates the
steps involved in the absorption mechanism. In the first step of the absorption
process, the solute diffuses from the bulk area of the gas phase to the gas-liquid
interface. In the second step, the solute transfers across the interface to the liquid
phase. In the third step, the solute diffuses into the bulk area of the liquid, making
room for additional gas molecules to be absorbed.
Bulk gas phase
Bulk liquid phase
Solute
Gas-liquid interface
Figure 4-1. Absorption process.
The purpose of analyzing these three steps is to determine which variables con-
trol the process. The most efficient system can be designed by knowing these
variables. It is assumed that once the solute arrives at the interface area, transfer
across it occurs instantaneously. This second step in the absorption mechanism is
extremely rapid. Therefore, it does not need to be considered when deriving
absorption efficiency equations. The rate of mass transfer (absorption) is dependent
upon the diffusion rate in either the gas phase or liquid phase.
Two terms used to describe the mass transfer rate are: gas phase controlled
absorption and liquid phase controlled absorption. Each mechanism depends on
the rate of diffusion in both phases and upon the solubility of the pollutant in the
liquid.
4-2
-------
The diffusion rate of a gaseous pollutant molecule through a gas is always faster
than its diffusion rate through a liquid because gas molecules in a gas are further
apart than are liquid molecules in a liquid. The pollutant molecule will move
faster through the gas phase than through the liquid phase since in a gas, obstruc-
tions such as other molecules art farther apart than in a liquid. However, the mass
transfer rate depends primarily upon the solubility -of the pollutant in the liquid.
Consider two gaseous pollutants, A and B, that are approximately equal in size.
Assume that A is not very soluble in the liquid while B is very soluble in the liquid.
Assume also that both A and B will move through the gas phase at the same rate.
The rates at which the pollutants will be absorbed in the liquid phase are different
because B is very soluble while A is not. A is absorbed very slowly and B is
absorbed very quickly.
The transfer rate of A (slowly absorbed) is greatly affected by its absorption and
diffusion in the liquid phase and is therefore referred to as liquid phase controlled.
The transfer rate of B (quickly absorbed) is also greatly affected by its rate of
movement through the liquid phase but is absorbed so rapidly that it doesn't
influence the overall transfer rate. The mass transfer rate for B depends only on
how fast it diffuses through the gas phase towards the gas-liquid interface and is
referred to as gas phase controlled.
The gas phase controlled systems absorb pollutants more readily than do the
liquid phase controlled systems. Thus, absorption systems used in the field of air
pollution control are usually designed to be gas phase controlled.
The previously mentioned three step mechanism of absorption occurs on two
levels: a micro or molecular level and a macro or bulk level. Molecular diffusion of
individual molecules occurs due to a concentration difference because all systems
try to achieve an equilibrium state. Molecules migrate from areas of high concen-
trations to areas of low concentrations. Macro or bulk diffusion occurs as a mass of
liquid or gas moves to or from the interface. In an absorber, bulk diffusion is
accomplished by turbulent mixing of the gas and liquid phases. To be efficient, an
absorber must provide both turbulent mixing of the gas and liquid phases and suf-
ficient residence time to allow pollutant molecules to be absorbed.
Solubility
A very important factor affecting the amount of a contaminant that can be
absorbed is the solubility of the contaminant. Solubility is a function of both the
temperature and to a lesser extent the pressure of the system. As temperature
increases, the amount of gas that can be absorbed by a liquid decreases. From the
ideal gas law: as temperature increases, the volume of a gas also increases;
therefore, at the higher temperature less gas is absorbed due to the increased
volume it occupies. Pressure affects the solubility of a gas in the opposite manner.
By increasing the pressure of a system the amount of gas absorbed generally
increases.
4-3
-------
The solubility of a specific gas in a given liquid is defined at a designated
temperature and pressure. Table 4-1 presents data on the solubility of SO2 gas in
water at 101 kPa or 1 atm and various temperatures. In determining solubility data,
the partial pressure (in mm Hg) is measured with the concentration (in grams of
solute per 100 g of liquid) of the solute in the liquid. The data in Table 4-1 was
taken from The International Critical Tables, a good source of information con-
cerning gas-liquid systems.
Table 4-1. Partial pressure of SO, in aqueous solution, mm Hg.
Grams
SO, per
100 g H,O
0.0
0.5
1.0
1.5
2.0
2.5
3.0
3.5
4.0
4.5
5.0
10°C
—
21
42
64
86
108
130
153
176
199
223
20 °C
—
29
59
90
123
157
191
227
264
300
338
30 °C
42
85
129
176
224
273
324
376
428
482
40°C
60
120
181
245
311
378
447
518
588
661
50 °C
_
83
164
247
333
421
511
603
698
793
—
60 °C
111
217
328
444
562
682
804
—
—
—
70 °C
144
281
426
581
739
897
—
Solubility data is obtained at equilibrium conditions. This involves putting
measured amounts of a gas and a liquid into a closed vessel and allowing it to sit
for a period of time. Eventually the amount of gas that is being absorbed into the
liquid will equal the amount which is coming out of the solution. At this point,
there is no net transfer of mass to either phase and the concentration of the gas in
both the gaseous and liquid phases remains constant. The gas-liquid system is at
equilibrium.
Equilibrium conditions are important in operating an absorption tower. If
equilibrium were to be reached in the actual operation of an absorption tower, the
collection efficiency would fall to zero at that point since no net mass transfer
would occur. The equilibrium concentration, therefore, limits the amount of solute
that can be removed by absorption. The most common method of analyzing
solubility data is to use an equilibrium diagram. An equilibrium diagram is a plot
of the mole fraction of solute in the liquid phase, denoted as x, versus the mole
fraction of solute in the gas phase, denoted as y. Equilibrium lines for the SO2 and
water system given in Table 4-1 are plotted in Figure 4-2. Figure 4-2 also illustrates
the temperature dependence of the absorption process. At a constant mole fraction
of solute in the gas (y), the mole fraction of SO2 in the liquid (x) increases as the
temperature decreases.
4-4
-------
0.6
.b 0-5
(S
O 0.4
t/3
.1 0.3
u
•3 0.2
0.1
0.002 0.006 0.010
Mole fraction of SOS in water
0.014
Figure 4-2. Equilibrium lines for SOt—H,O system.
Under certain conditions, Henry's Law may also be used to express equilibrium
solubility of gas-liquid systems. Henry's Law is expressed as:
(Eq. 4-1)
P* = .
Where: p* = partial pressure of solute at equilibrium, Pa
x = mole fraction of solute in the liquid
.')(' = Henry's Law constant, Pa/mole fraction
Henry's Law can be written in a more useful form by dividing both sides of Equa-
tion 4-1 by the total pressure, PT, of the system. The left side of the equation
becomes the partial pressure divided by the total pressure which equals the mole
fraction in the gas phase, y*. Equation 4-1 now becomes:
(Eq. 4-2)
y* =
Where: y* = mole fraction in gas phase in equilibrium with liquid
,'j(" = Henry's Law constant, mole fraction in vapor/mole fraction in
liquid
Note: ,'J(" is now dependent on the total pressure
4-5
-------
Equation 4-2 is the equation of a straight line, where the slope (m) is equal to Jf.
Henry's Law can be used to predict solubility only when the equilibrium line is
straight. Equilibrium lines are usually straight when the solute concentrations are
very dilute. In air pollution control applications this is usually the case. For exam-
ple, an exhaust stream that contains a 1000 ppm SO2 concentration corresponds to
a mole fraction of SO2 in the gas phase of only 0.001. Figure 4-2 demonstrates that
the equilibrium lines are still straight at this low concentration of SO2.
Another restriction on using Henry's Law is that it does not hold true for gases
that react or dissociate upon dissolution. If this happens, the gas no longer exists
as a simple molecule. For example, scrubbing HF or HC1 gases with water causes
both compounds to dissociate in solution. In these cases, the equilibrium lines are
curved rather than straight. Data on systems that exhibit curved equilibrium lines
must be obtained from experiments.
Henry's Law constants for the solubility of several gases in water are listed in
Table 4-2. The units of Henry's Law constants are atm/mole fraction. The smaller
the constant, the more soluble the gas. Table 4-2 demonstrates that SO2 is approx-
imately 100 times more soluble in water than CO2 is in water. The following exam-
ple illustrates how to develop an equilibrium diagram from solubility data.
Table 4-2. Henry's Law constants for gases in H8O*.
Gas
N2
CO
HjS
02
NO
CO2
sof
20°C
80.4
53.6
48.3
40.1
26.4
1.42
0.014
30°C
92.4
62.0
60.9
47.5
31.0
1.86
0.016
* Expressed in J/'x 10'3, atm/mole fraction.
Source: Marchello, 1976.
4-6
-------
Example 4-1
Given the following data for the solubility of SO2 in pure water at 303 K (30 °C)
and 101.3 kPa (760 mm Hg), plot the equilibrium diagram and determine if
Henry's Law applies.
Equilibrium data
(g of SO, per 100 g of H,O)
0.5
1.0
1.5
2.0
2.5
3.0
(partial pressure of SO*)
6 kPa (42 mm Hg)
11.6kPa (85 mm Hg)
18.3 kPa (129 mm Hg)
24.3 kPa (176 mm Hg)
30.0 kPa (224 mm Hg)
36.4 kPa (273 mm Hg)
Solution:
1. The data must first be converted to mole fraction units.
The mole fraction in the gas phase, y*, is obtained by dividing the partial
pressure of SO2 by the total pressure of the system:
= 0.06
= P*°z =
Pr 101.3kPa
The mole fraction in the liquid phase, x, is obtained by dividing the moles of
SO2 in solution by the total moles of liquid:
moles SO2 in solution
x =
moles SO2 in solution + moles of H2O
moles of SO2 in solution = c.so2/64 g SO2 per mole
moles of H2O= 100 gm of H2O/18 g H2O per mole=5.55
0.5
Cyo,/64
64
x =
Cso,/64+5.55
0.5
64
-0.0014
+ 5.55
4-7
-------
2. Completing the following for Example 4-1:
g of SO,
Ccn. a —
1 lOOgH.O
0.5
1.0
1.5
2.0
2.5
3.0
P*
(kPa)
6.0
11.6
18.3
24.3
30.0
36.4
y*=p/101.3
0.060
0.115
0.180
0.239
0.298
0.359
Csot/64
Cjo/64+5.55
0.0014
0.0028
0.0042
0.0056
0.0070
0.0084
The above data is plotted in Figure 4-3. Henry's Law applies in the given con-
centration range with Henry's Law constant equal to 42.7 mole fraction SO2 in
air/mole fraction SO2 in water.
o
c/)
o
_
~O
>^ 10 -
0.004 0.008
x, mole fraction of SOj in water
0.012
Figure 4-3. Equilibrium diagram for SO,—H,O system.
Absorption Design Theory
The first step in designing an air pollution control device is to develop a
mathematical expression describing the observed phenomenon. A valid
mathematical expression describing absorber performance makes it possible to:
determine the proper absorber size for a given set of conditions, and predict how a
change in operating conditions affects absorber performance. There are a number
of theories or models which attempt to analytically describe the absorption
4-8
-------
mechanism. However, in practice, none of these analytical expressions can solely be
used for design calculations. Experimental or empirical data must also be used to
obtain reliable results.
The basic model for describing the absorption process is the two film or double
resistance theory which was first proposed by Whitman in 1923. The model starts
with the three step mechanism of absorption, previously discussed and illustrated in
Figure 4-1. From this mechanism, the rate of mass transfer was shown to be depen-
dent on the rate of diffusion of a molecule in either the gas or liquid phase. The
two film model starts by assuming that the gas and liquid phases are in turbulent
contact with each other, separated by an interface area where they meet. This
assumption may be correct; but no mathematical expressions adequately describe
the diffusion (transport) of a molecule through both phases in turbulent motion.
Therefore, the model proposes that a mass transfer zone exists to include a small
portion (film) of the gas and liquid phases on either side of the interface. The mass
transfer zone is comprised of two films, a gas film and a liquid film on their
respective sides of the interface. These films are assumed to flow in a laminar or
streamline motion, for which mathematical expressions that describe diffusion do
exist. This concept of the two film theory is illustrated in Figure 4-4.
Bulk gas phase
Bulk liquid phase
Partial pressure
driving force
Concentration
driving force
Gas Liquid
film film
Figure 4-4. Visualization of two film theory.
4-9
-------
According to the two film theory, for a molecule of substance A to be absorbed
it must proceed through a series of five steps. The molecule must diffuse:
1. from the bulk gas phase to the gas film,
2. through the gas film
3. across the interface,
4. through the liquid film,
5. and finally mix into the bulk liquid.
The theory assumes that there is complete mixing in both the gas and liquid bulk
phases and that the interface is at equilibrium with respect to molecules trans-
ferring in or out. This implies that all resistance to diffusion occurs when the
molecule is diffusing through the gas and liquid films, hence the name double
resistance theory. The concentration in the gas phase changes from p^c in the bulk
gas to PAT at the interface.
A gas concentration is expressed by its partial pressure. Similarly, the concentra-
tion in the liquid changes from CAI at the interface to cAl in the bulk phase as mass
transfer occurs. The rate of mass transfer is then equal to the amount of molecule
A transferred times the resistance molecule A encounters in diffusing through the
films:
- 4-3) NA
=mass transfer coefficient for liquid film, g mole/hr«m2«Pa
(Ib mole/hr»ft2»atm)
The mass transfer coefficients, k, and kf, represent the flow resistance the solute
encounters in diffusing through each film respectively (Figure 4-5). An analogy is
the resistance electricity encounters as it flows through a circuit.
4-10
-------
Total resistance
Figure 4-5. Resistance to motion encountered by a molecule being absorbed.
Equations 4-3 and 4-4 define the general case of absorption and are applicable
to both curved and straight equilibrium lines. In practice, Equations 4-3 and 4-4
are difficult to use, since it is impossible to measure the interface concentrations p
and CAT. The interface is a fictitious quantity used in the model to represent an
observed phenomenon. The interface concentrations can be avoided by defining
the mass transfer system at equilibrium conditions and combining the individual
film resistances into an overall resistance. If the equilibrium line is straight, the
rate of absorption is given by:
(Eq. 4-5) Nx = KoG(p*c-p;0
(Eq. 4-6) NX = K0i (cj? - CAL)
Where: pi = equilibrium partial pressure of solute A at operating conditions
cjf = equilibrium concentration of solute A at operating conditions
Koo3 overall mass transfer coefficient based on gas phase,
g mole/hr«m2»Pa (Ib mole/hr«ft2»atm)
KOI = overall mass transfer coefficient based on liquid phase,
g mole/hr»m2»Pa (Ib mole/hr»ft2»atm)
4-11
-------
An important fact to note concerning Equations 4-5 and 4-6 is that they impose
an upper limit on the amount of solute that can be absorbed. The rate of mass
transfer is dependent on the concentration departure from equilibrium in either
the gas (PAG- p3) or liquid (of - cAi) phase. The larger these concentration
differences are, the greater the rate of mass transfer. If equilibrium is ever reached
(PAG = P$ or CAL = cjj) absorption stops and no net transfer occurs. Thus, the
equilibrium concentrations determine the maximum amount of solute that is
absorbed.
At equilibrium the overall mass transfer coefficients are related to the individual
mass transfer coefficients by:
(Eq. 4-7) J_ = J_ + £
k, k£
(Eq. 4-8) — !— = — + — !—
4 Kor k£ .'M'\
,'JC is Henry's Law constant (the slope of the equilibrium). Equations 4-7 and 4-8
are useful in determining which phase controls the rate of absorption. From Equa-
tion 4-7, if J£ is very small (which means the gas is very soluble in the liquid) then
Koo — kg, and absorption is said to be gas film controlled. The major resistance to
mass transfer is in the gas phase. Conversely, if a gas has limited solubility 3C is
large, and Equation 4-8 reduces to KO£ «kp. The mass transfer rate is liquid film
controlled and depends on the solute's dispersion rate in the liquid phase. Most
systems in the air pollution control field are gas phase controlled since the liquid is
chosen so that the solute will have a high degree of solubility.
The discussion so far has been based on the two film theory of absorption. There
are other theories which offer different descriptions of gas molecule movement
from the gas to the liquid phase. For these theories, the mass transfer rate equation
does not differ from that of the two film model. The difference lies in the way they
predict the mass transfer coefficient. It has been shown that the rate of mass
transfer is dependent on a concentration difference multiplied by a resistance
factor. Like most theories describing how something functions, absorption theories
provide a basic understanding of the process, but due to the complexities of "real
life" operations it is difficult to apply them directly. Concentrations can easily be
determined from operating (c and p) and equilibrium (cjj and p*) data of the
system. Mass transfer coefficients are very difficult to determine from theory.
Theoretically predicted values of the individual mass transfer coefficients (kg and
kf) based on the two film theory, do not correlate well with observed values.
Overall mass transfer coefficients are more easily determined from experimental or
operational data. However, the overall coefficients apply only when the
equilibrium line is straight.
Mass transfer coefficients are often expressed by the symbols K0ca, kfa, etc.
where "a" represents the surface area available for absorption per unit volume of
the column. This allows for easy determination of the column area required to
accomplish the desired separation. These mass transfer coefficients are developed
from experimental data and usually reported in one of two ways: as an empirical
relationship based on a function of the liquid flow or gas flow or slope of the
4-12
-------
equilibrium line; or correlated in terms of a dimensionless number, usually either
the Reynolds or Schmit Number. Figure 4-6 compares the effect on the mass
transfer coefficient for SO2 in water using two types of packing materials (Perry,
1973). Packing A is one-inch rings and packing B is three-inch spiral tiles. Similar
figures are used extensively to compare different absorbers or similar absorbers with
varying operating conditions.
20
M
<£
\
f
•
£*
•
I-
^ 6
"o
E
o 4
IS
J
1 1 ' 1 1 II'
,^ — •• " "" """"
^ — ^~
" ^>^*^^~
- \^
— -
1 1 1 1 1 111
20 40 80 100 200 400 500 100
Where:
Packing A= 1 in. rings
Packing B = 3 in. spiral tiles
G', Ibs/hr-ft2
Source: Perry 1973.
Figure 4-6. Comparison of overall absorption coefficient for SOj in water.
Although the science of absorption is considerably developed, much of the work
in practical design situations is empirical in nature. The following section will
apply the principles discussed, to the design of gas absorption equipment. Emphasis
has been placed on presenting information that can be used to estimate absorber
size and liquid flow rate.
Design Procedures
Effective gas absorption depends on intimate contact between the gas and liquid
phases. This requires a high degree of turbulence between phases. However, high
turbulent flow can result in extremely short contact time between the two phases,
limiting absorption efficiency. The two most common absorbers are plate absorbers
and packed tower absorbers. Plate absorbers are vertical towers with one or more
plates mounted horizontally inside (see Figure 4-28). Gas enters the bottom of the
tower and flows upward through openings in the plates. Liquid is distributed across
the plates and mixing occurs as the gas bubbles up through the plates. Packed
towers are either horizontal or vertical columns with liquid sprays and packing
material (see Figure 4-23). Liquid introduced by sprays or weirs coats the inert
4-13
-------
packing material that provides a large surface area for continuous gas-liquid con-
tact. Both of these devices are used extensively to control gaseous pollutants.
Absorber design calculations presented in this chapter will focus on these two
devices.
Numerous procedures are used to design an absorption system. These procedures
range in difficulty and cost from shortcut "rules of thumb" equations to in-depth
design procedures based on pilot plant data. Procedures presented here will be
based on the shortcut "rules of thumb". These procedures allow for a quick review
of proposed system designs and an easy estimation of process change effects on
absorber operation.
To design an absorption system, certain parameters are set by either operating
conditions or regulations. The gas stream to be treated is usually the exhaust from
a process in the plant. Therefore, the volume, temperature, and composition of the
gas stream are given parameters. The outlet composition of the contaminant is set
by the emission standard which must be met. The temperature and inlet com-
position of the absorbing liquid are also usually known. The main unknowns in
designing the absorption system are: the flow rate of liquid required; the diameter
of the vessel needed to accommodate the gas and liquid flow; and the height of
absorber required to achieve the needed removal. Procedures for estimating these
three unknowns will be discussed in the following sections.
Material Balance
In designing or reviewing the design of an absorption control system, the first task
is to determine the flow rates and composition of each stream entering the system.
From the law of conservation of mass, the material entering a process must either
accumulate or exit. In other words, "what comes in must go out". A material
balance is used to help determine flow rates and compositions of individual
streams. Figure 4-7 illustrates a typical countercurrent flow absorber in which a
material balance is drawn. The solute is used as the "material" in the material
balance.
4-14
-------
Y,
Figure 4-7. Material balance for countercurrent flow absorber.
The following procedure to set up a material balance and determine the liquid
flow rate will focus on a countercurrent gas-liquid flow pattern. This is the most
common flow pattern used to achieve high efficiency gas absorption. For cocurrent
flow, only a slight modification (to inlet and outlet gas flows) of this procedure is
required. Equations for crosscurrent flows are very complicated as they involve a
gradient pattern that changes in two directions. They will not be presented here.
The terms in Figure 4-7 are defined as:
X mole fraction of solute in pure liquid
Y mole fraction of solute in inert gas
Lm liquid flow rate, g mole/hr
Gm gas flow rate, g mole/hr
Engineering design work is usually done on a solute free basis (X, Y), to make the
material balance calculations easier. The solute free basis is defined as:
(Eq. 4-9)
(Eq. 4-10)
Y=
V —
1-y
x
1-x
4-15
-------
In air pollution control systems the percent of pollutant transferred, y and x, is
generally small. Therefore, from Equations 4-9 and 4-10 Y = y and X==x. In this
chapter, it is assumed that X and Y are always equal to x and y respectively. If
y (inlet gas concentration) ever becomes larger than a few percent by volume this
assumption is invalid and will cause errors in the material balance calculations.
An overall mass balance across the absorber in Figure 4-7 yields:
(Eq. 4-11) Ib mole in = lb mole out
Gm(in) + Lm(in) = Gm(out) + Lm(out)
For convenience the top of the absorber is labeled as point 2 and the bottom as
point 1. This changes Equation 4-11 to:
(Eq. 4-12) Gml + Lm2 = Gm2 + Lml
In this same manner, a material balance for the contaminant to be removed is
obtained:
(Eq. 4-13) GmlY, + Lm2X2 = Gm2Y2 + LmlX,
Equation 4-13 can be simplified by assuming that as the gas and liquid streams
flow through the absorber their total mass does not change appreciably (i.e.
Gml = Gm2 and Lmt = Lm2). This is justifiable for most air pollution control systems
since the mass flow rate of contaminant is very small compared to the liquid and
gas mass flow rates. For example, a 10,000 cfm exhaust stream containing 1000
ppm SO2 would be only 0.1% SO2 by volume or 1.0 cfm. If the scrubber were
100% efficient the gas mass flow rate would change from 10,000 cfm at Gmi to
9,999 cfm at Gm2. The transfer of a quantity this small is negligible in an overall
material balance. Therefore, Equation 4-13 can be reduced to:
(Eq. 4-14) Gm(Y, - Y2) = Lm(X2 - X.)
and rearranging
(Eq. 4-15) Y,-Y2= t=(X,-X.)
Gm
Equation 4-15 is the equation of a straight line. When this line is plotted on an
equilibrium diagram it is referred to as an operating line. This line defines
operating conditions within the absorber; what is going in and what is coming out.
An equilibrium diagram with a typical operating line plotted on it is shown in
Figure 4-8. The slope of the operating line is the liquid mass flow rate divided by
the gas mass flow rate which is the liquid-to-gas ratio or Lm/Gm. The liquid-to-gas
ratio is used extensively when describing or comparing absorption systems. Deter-
mination of the liquid-to-gas ratio is discussed in the next section.
4-16
-------
Slope of operating line =
Lm
Gm
X
Figure 4-8. Typical operating line diagram.
Determining the Liquid Requirement
In the design of most absorption columns the quantity of gas to be treated (Gm)
and the inlet solute concentration (Y]) are set by process conditions. Minimum
acceptable standards specify the outlet solute concentration (Y2). The composition
of the liquid into the absorber (X2) is also generally known or can be assumed to
be zero if there is no recycling of liquid. By plotting this data on an equilibrium
diagram, the minimum amount of liquid required to achieve the required outlet
concentration (Y2) can be determined.
4-17
-------
Figure 4-9a is a typical equilibrium diagram with operating points plotted for a
countercurrent flow adsorber. At the minimum liquid rate, the inlet gas concentra-
tion of solute (Yt) is in equilibrium with the outlet liquid concentration of solute
(Xmax). The liquid leaving the absorber is saturated with solute and can no longer
dissolve any more solute unless additional liquid is added. This condition is
represented by point B on the equilibrium curve.
The slope of the line drawn between point A and point B represents the
operating conditions at minimum flow rate in Figure 4-9b. Note how the driving
force decreases to zero at point B. The slope of line AB is (Lm/Gm)min, and may
be determined graphically or from the equation for a straight line. By knowing the
slope of the minimum operating line, the minimum liquid rate can easily be deter-
mined by substituting in the known gas flow rate. This procedure is illustrated in
Example 4-2.
Determining the minimum liquid flow rate, (Lm/Gm)min, is important since
absorber operation is usually specified as some factor of it. Generally, liquid flow
rates are specified at 25 to 100% greater than the required minimum. Typical
absorber operation would be at 50% greater than the minimum liquid flow rate
(i.e. 1.5 times the minimum liquid-to-gas ratio). Setting the liquid rate in this way
assumes that the gas flow rate set by the process does not change appreciably. Line
AC in Figure 4-9c is drawn at a slope of 1.5 times the minimum Lm/Gm. Line AC
is referred to as the operating line since it describes absorber operating conditions.
4-18
------- An error occurred while trying to OCR this image.
-------
The following example problem illustrates how to compute the minimum liquid
rate required to achieve a desired removal efficiency.
Example 4-2
Using the data and results from Example 4-1, compute the minimum liquid rate of
pure water required to remove 90% of the SO2 from a gas stream of 84.9 mVmin
(3000 acfm) containing 3% SO2 by volume. The temperature is 293 K and the
pressure is 101.3 kPa.
Solution:
1. First, sketch and label a drawing of the system.
Y2 = 0.003
L=?
Xz = 0
^>d
Q.= 84.9 m'/min
Y, = 0.03
X, =
by volume =0.03
Y2 = 90% reduction from YI or only 10% of YI; therefore
¥, = (0.10X0. 03) = 0.003
2. At the minimum liquid rate Y, and Xi will be in equilibrium; the liquid will be
saturated with SO2.
At equilibrium:
W- 4.9 7 mole fraction SO2 in air
.Jt — 4Z.7 - (from Example 4-1)
mole fraction SO2 in water
0.03 = 42.7 X,
Xi = 0.000703 mole fraction
3. The minimum liquid-to-gas ratio from Equation 4-15 is:
V,-Y,= i?v
Gm
u = Y!-YZ
cm x, - x2
4-20
-------
Lm _ 0.03-0.003 _ ofi 4 g mole water
Gm ~ 0.000703-0 ' g mole of air
4. To compute the minimum required liquid flow rate:
First, convert m3 of air to g moles.
At 0°C and 101.3 kPa there are 0.0224 mVg mole (359 ft3/ mole) of an ideal gas.
At 30°C: 0.0224
m3
g mole \273 K
293 K
= 0.024 ms/g mole
Gm = 84.9
m3
g mole air \ _ QeQO g mole air
min \ 0.024 m3
mm
Lm „„ . g mole water . . ,. .
—- =38.4 -2 at minimum conditions
Gm g mole air
Lm = (38.4)(3538)= 136.0 kg mole water min.
In mass units:
136 kg mole / 18 kg \ nAAQ kg ,c _ , , . N
L= 2 §_ =2448—^-(647 gal/mm)
min \kg mole/ min
5. Figure 4-10 illustrates the graphical solution to this problem.
Multiply the slope of the minimum operating line times 1.5 to get the
slope of the actual operating line (line AC).
38.4x1.5 = 57.6
Minimum operating line
0.0002 0.0004 0.0006
X, mole fraction of SO2 in water
Figure 4-10. Solution to Example 4-2.
0.0008
4-21
-------
Sizing the Absorber
Packed Column Diameter
The main parameter which affects the size of a packed column is the gas velocity
at which liquid droplets become entrained in the exiting gas stream. Consider a
packed column operating at set gas and liquid flow rates. By decreasing the
diameter of the column, the gas flow rate (m/s or ft/sec) through the column will
increase. If the gas flow rate through the column is gradually increased (by using
smaller and smaller diameter columns) a point will be reached where the liquid
flowing down over the packing begins to be held in the void spaces between the
packing. This gas to liquid flow ratio is termed the loading point. The pressure
drop over the column begins to increase and the degree of mixing between the
phases decreases. A further increase in gas velocity will cause the liquid to com-
pletely fill the void spaces in the packing. The liquid forms a layer over the top of
the packing and no more liquid can flow down through the tower. The pressure
drop increases substantially and mixing between the phases is minimal. This condi-
tion is referred to as flooding and the gas velocity at which it occurs is the flooding
velocity. Using an extremely large diameter tower would eliminate this problem.
However, as the diameter increases the cost of the tower increases.
Normal practice is to size a packed column diameter to operate at a certain per-
cent of the flooding velocity. A typical operating range for the gas velocity through
the columns is 50 to 75% of the flooding velocity. It is assumed that by operating
in this range the gas velocity will also be below the loading point.
A common and relatively simple procedure to estimate the flooding velocity (thus
setting a minimum column diameter) is to use a generalized flooding and pressure
drop correlation. One version of the flooding and pressure drop relationship in a
packed tower is the Sherwood Correlation shown in Figure 4-11 (EPA, 1972). This
correlation is based on the physical properties of the gas and liquid streams and
tower packing characteristics. The procedure to determine the tower diameter is:
1. Calculate the value of the abscissa.
(Eq. 4-16) -
Where: L and G= mass flow rates: any consistant set of units may be used as
long as the abscissa is dimensionless
QS = density of the gas stream
Qf= density of the absorbing liquid
2. From the point calculated in Equation 4-16 proceed up the graph to the
flooding line and read the ordinate e.
3. Rearrange the equation of the ordinate and solve for G'.
r T
(Eq. 4-17) G' = (e)(e*)(el)(gc)
L F0u,o-« J
4-22
-------
Where: F= packing factor given in Table 4-3 for different types of packing
(Bhatia, 1977)
4> = ratio of specific gravity of the scrubbing liquid to that of water
JJL^= viscosity of liquid (for water = 0.8 centipoise)
G' = mass flow rate of gas per unit cross sectional area of column at
flooding, g/s»m2 (lb/s«ft2)
Qg= density of the absorbing liquid, kg/m3 (lb/ft3)
Qg = density of the gas stream, kg/m3 (lb/ft3)
gc = gravitational constant = 9.82 m/s2 (32.2 ft/s2)
4. G' at operating is a fraction of G' at flooding.
(Eq. 4-18) G operating— (i)(G flooding)
5. The cross sectional area of column A is calculated from:
(Eq. 4-19) A=-7r,
vjr operating
6. The diameter of the column is obtained from:
(Eq. 4-20)
/4A\°-S
d,= ( — =1.13 A05
7T/
i
oi
0.5
0.2
0.1
0.05
0.02
0.01
0.005
0.002
0.001
Pressure drop,
m H2O/m packing (in. H2O/ft packing)
•0-0416(0.5)
•0.0208(0.25)
•0.00808(0.10)
=-0.00416(0.05)
0.01 0.02 0.05 0.1 0.2 0.5 1
L
G
(dimensionless)
5 10
Figure 4-11. Generalized flooding and pressure drop correlation.
4-23
------- An error occurred while trying to OCR this image.
-------
Example 4-3
For the scrubber in Example 4-2, determine the column diameter if the operating
liquid rate is 1.5 times the minimum. The gas velocity should be no greater than 75%
of the flooding velocity and the packing material is two-inch ceramic Intalox saddles.
Solution:
From Example 4-2:
Gm = 3538 g mole/min
L(m,«imMm) = 2448 kg/min
Convert gas molar flow to a mass flow, assuming molecular weight of the gas to be
29 kg/mole.
kg/min
min mole
Adjusting the liquid flow to 1.5 times the minimum:
L=1.5 (2448) = 3672 kg/min
The densities of air and water at 30 °C are:
Q£= 1000 kg/m3
<>,= 1.17kg/m»
1. Now using the relationship in Equation 4-16, compute the abscissa.
\102.6/\1000/
2. From Figure 4-11, proceed up to the flooding line from 1.22. The ordinate e is
0.019.
3. Use Equation 4-17 to solve for the superficial flooding velocity, G'.
A superficial velocity is a flow rate per unit cross sectional area.
_
Vjr —
For water, = 1.0 and /*£= 0.0008 Pa«s
From Table 4-3 for two-inch Intalox saddles, F=40 ftVft3 or 131 mVm3.
gc = 9.82 m/s2
r(0.019)(1.17)(1000)(9.82) °5
(1)(131)(0.0008)02
= 2.63 kg/m2»s at flooding
1 °
J
4-25
-------
4. The superficial gas velocity at operating is obtained from Equation 4-18.
G'op.ra(,-ng = fG'/(ood.ng = (0.75)(2.63)= 1.97 kg/m2.s
5. From Equation 4-19, the cross sectional area of the tower is:
A- G
ting
Gi
operati
_ (102.6 kg/min)(min/60s) _ Q 3? m?
1.97 kg/mz-s
6. The diameter of the tower, from Equation 4-20 is:
/4A\»
7T/
_ f(4)(0.87 m2)] °5
~L * J
= 1.05 m or at least 1.1 m (3.5 ft)
Figure 4-11 may also be used to estimate the pressure drop of the tower once
^-* operating IS Set.
This is done by plugging G' back into the equation to compute e.
_(1.97)2(1)(131)(0.0008)02
(1.17X1000)(9.82)
= 0.0106
The abscissa remains unchanged and equals 1.22.
The pressure drop through the column is the point at which these two lines cross.
From Figure 4-11:
A _
0.0416 meter of H2O /0.5 inches of H2O
meter of packing \ feet of packing
Packed Column Height
The height of a packed column refers to the depth of packing material needed to
accomplish the required removal efficiency. The more difficult the separation, the
larger the packing height required. For example, a much larger packing height
would be required to remove SO2 than to remove Cl from an exhaust stream using
water as the absorbent. This is because Cl is more soluble in water than SO2.
Determining the proper height of packing is important since it affects both the rate
and efficiency of absorption.
4-26
-------
A number of theoretical equations are used to predict the required packing
height. These equations are based on diffusion principles. Depending on which
phase is controlling the absorption process, either Equation 4-5 or 4-6 is used as the
starting point to derive an equation to predict column height. A material balance
is then set up over a small differential section of the column. Derivations of the
equations used can be found in a number of chemical engineering texts or other
books dealing with mass transfer, and will not be covered here.
The general form of the design equation for a gas phase controlled resistance is
given by:
Y,
(Eq.4-21) Z= -£1 1
K,aP YZ (1-Y)(Y-Y*)
Where: Z = height of packing, m
a = interfacial contact area, mz
P = pressure of the system, kPa
In analyzing Equation 4-21, the term G'/KgaP has the dimension of meters and
is defined as the height of a transfer unit. The term inside the integral is dimen-
sionless and represents the number of transfer units needed to make up the total
packing height. Using the concept of transfer units, Equation 4-21 can be
simplified to:
(Eq. 4-22) Z = HTU x NTU
Where: HTU = height of a transfer unit, m
NTU = number of transfer units
The concept of a transfer unit comes from the operation of plate columns.
Discrete stages (plates) of separation occur in plate columns. These stages can be
visualized as a transfer unit with the number and height of each giving the total
tower height. Although packed columns operate as one continuous separation
process, in design terminology it is treated as if it were broken into discrete sections
(height of a transfer unit). The number and the height of a transfer unit are based
on either the gas or liquid phase. Equation 4-22 now becomes:
(Eq. 4-23) Z - NocHoc = N0£H
0£ot
Where: NQC^ number of transfer units based on overall gas film coefficient
NOZ = number of transfer units based on overall liquid film coefficient
HOG = height of a transfer unit based on overall gas film coefficient, m
H0£ = height of a transfer unit based on an overall liquid film coefficient, m
4-27
-------
Values for the height of a transfer unit used in designing absorption systems are
usually obtained from experimental data. To ensure greatest accuracy, vendors of
absorption equipment normally perform pilot plant studies to determine the height
of a transfer unit. For common absorption systems, such as NHs-water, manufac-
turers have developed graphs which can be used to estimate the height of a transfer
unit. These graphs do not provide the accuracy of pilot plant data, but they are
less expensive and easier to use. Figure 4-12 gives a typical example of these graphs
for an ammonia-water system. In this figure, the superficial gas flow rate is plotted
versus the Hoc with the superficial liquid rate as a parameter. It is also common to
plot liquid rate versus the Hoc and have the gas rate as a parameter. Additional
information on other gas-liquid systems can be found in Chemical Engineers'
Handbook (Perry, 1973). In applying these data it is important that the process
conditions be similar to the conditions at which the HTU was measured.
3.6
2.8
§ 2.0
1.2
0.4
G' = 500 Ib/hr-ft2
I
I
500
Where:
o = 11/£ in. Raschig rings
A = 1 in. Tellerettes
1000
L', lb/hr.ft2
1500
2000
Figure 4-12. Column packing comparison for
ammonia and water system.
When no experimental data are available or if only a preliminary estimate of
absorber efficiency is needed, there are generalized correlations available to predict
the height of a transfer unit. The correlations for predicting the HOC or the H0i
are empirical in nature and are a function of: type of packing; liquid and gas flow
rates; concentration and solubility of the contaminant; liquid properties; and
system temperature. These correlations can be found in engineering texts such as
Chemical Engineers' Handbook (Perry, 1973), EPA (1972), or Treybal (1968). For
most applications, the height of a transfer unit ranges between 0.3 to 1.2 m (1 to
4 ft) (Calvert, 1977). As a rough estimate, 0.6 m (2.0 ft) can be used.
4-28
-------
The number of transfer units, NTU, can be obtained experimentally or
calculated from a variety of methods. For the'case where the solute concentration is
very low and the equilibrium line is straight, Equation 4-24 can be used to deter-
mine the number of transfer units (N0c) based on the gas phase resistance. Equa-
tion 4-24 can be derived from the integral portion of Equation 4-21.
(Eq. 4-24)
In
Y
Y
i
2
-mX
-mX
2
2
('*
mGmx
U
\ mG
\ _(_
1 u
m
1-
mGn
Where: m = slope of equilibrium line
Gm = molar flow rate of gas, kg mole/hr
Lm = molar flow rate of liquid, kg mole/hr
Xz = mole fraction of solute entering the column
Yj = mole fraction of solute in entering gas
Y2 = mole fraction of solute in exiting gas
Equation 4-24 may be solved directly or graphically by using the Colburn
Diagram which is presented in Figure 4-13. The Colburn Diagram is a plot of the
NOG versus ln[Yj - mX2/Y2 - mX2] at various values of (mGm/Lm). The term
(mGm/Lm) is referred to as the absorption factor. Figure 4-13 is used by first com-
puting the value of ln[Yi — mX2/Y2 - mX2], reading up the graph to the line cor-
responding to (mGm/Lm), and then reading across to obtain the NOG-
4-29
-------
C
3
I
o
I*
-------
The number of transfer units only depends on the inlet and outlet concentration of
the solute (contaminant). For example, if the conditions of Equation 4-25 are met,
to achieve 90% removal of any pollutant requires 2.3 transfer units. Equation 4-25
applies only when the equilibrium line is straight and approaches zero (for very
soluble or reactive gases). The following example illustrates the use of Equation 4-24.
Example 4-4
From pilot plant studies of the absorption system in Example 4-2, it was deter-
mined that the HOG for the SO2-water system is 0.829 m (2.72 ft). Calculate the
total height of packing required to achieve the 90% removal. The following data
were taken from the previous examples:
m = 427kgmoleH20
kg mole air
mm
Lm = 3672-^g-x kfe " =204
min 18 kg mm
X2 = 0 (no recycle liquid)
Y2 = 0.003
Solution:
Compute the N0c from Equation 4-24:
.JY.-mX, /_mC.\ mG.1
NT _ L Y2 ~ mX2 \ L"- / L- J
mG,,
i f/0-03
In
LVO.OOS
03V (42.7)(3.5)\ ( (42.7)(3.5)
OOS/V 204 / 204
_ (42.7)(3.5)
204
= 4.58
The total packing height is thus:
Z = HOG X NOG
Z = (0.829)(4.58) = 3.79 m of packing height
4-31
-------
Plate Tower
In a plate tower the scrubbing liquid enters at the top of the tower, passes over the
top plate, and then down over each lower plate until the liquid reaches the
bottom. Absorption occurs as the gas, which enters at the bottom, passes up
through the plate and contacts the liquid. In a plate tower, absorption occurs in a
stepwise or stage process. The operation of plate towers is discussed in greater
detail in the equipment section of this manual.
There are various accepted procedures available to size a plate tower. Detailed
summaries can be found in Chemical Engineers' Handbook (Perry, 1973), McCabe
and Smith (1956), and Theodore and Buonicore (1975). The following discussion
presents a simplified method for sizing or reviewing the design of a plate tower.
Plate Tower Diameter
The minimum diameter of a single-pass plate tower is determined by using the gas
velocity through the tower. If the gas velocity is too fast, liquid droplets are
entrained, causing a condition known as priming. Priming occurs when the gas
velocity through the tower is so fast that it causes liquid on one tray to foam and
then rise to the tray above. Priming reduces absorber efficiency by inhibiting gas
and liquid contact. For the purpose of determining tower diameter, priming in a
plate tower is analogous to the flooding point in a packed tower. It determines the
minimum acceptable diameter. The actual diameter should be larger.
The smallest allowable diameter for a plate tower is expressed by:
(Eq. 4-26)
Where: Q= volumetric gas flow, mVhr
\£ = empirical correlation, m° 25 hr° 25/kg° "
Qg = gas density, kg/m3
The term i/- is an empirical correlation and is a function of both the tray spacing
and the densities of the gas and liquid streams. Values for i/' in Table 4-4 are for a
tray spacing of 61 cm (24 in.) and a liquid specific gravity of 1.05 (EPA, 1972). If
the specific gravity of a liquid varies significantly from 1.05, the values for \l/ in
Table 4-4 cannot be used.
Table 4-4. Empirical constants for Equation 4-26.
Tray
Bubble cap
Sieve
Valve
Metric \l/a
0.0162
0.0140
0.0125
Engineering i//6
0.1386
0.1198
0.1069
"Metric !/• is expressed in m° "hr° Vkg° ", for use with
expressed in mVhr, and QS expressed in kg/m3.
'English \l> is expressed in ft0 25min°-5/lb° ", for use with
in cfm, and Q, expressed in lb/ft3.
Source: EPA, 1972.
4-32
-------
Depending on operating conditions, trays are spaced at a minimum distance
between plates to allow the gas and liquid phases to separate before reaching the
plate above. Trays should be spaced to allow for easy maintenance and cleaning.
Trays are normally spaced 45 to 70 cm (18 to 28 inches) apart. In using Table 4-4
for a tray spacing different than 61 cm, a correction factor must be used. Figure
4-14 is used to determine the correction factor which is multiplied times the
estimated diameter. Example 4-5 illustrates how to estimate the minimum diameter
of a plate tower.
Tray spacing, inches
1.5
15
I
18
I
21
I
TJ
<2
O
W
U
u
u
O
U
1.3
1.1
1.05
0.9
24
]_
27
I
30
I
0.3 0.4 0.53 0.6
Tray spacing, meters
0.8
Figure 4-14. Tray spacing correction factor.
4-33
-------
Example 4-5
For the condition described in Example 4-2, determine the minimum acceptable
diameter if the scrubber is a bubble cap tray tower. The trays are spaced 0.53 m
(21 inches) apart.
Solution:
From Example 4-2, the following information is obtained:
Gas flow rate= Q= 84.9 mVmin
Density = o,= 1.17 kg/m3
From Table 4-4 for a bubble cap tray:
^=0.0162 m°-»hr025/kg025
Before Equation 4-26 can be used, Q must be converted to mVhr.
= 84.9 X= 5094 m'/hr
mm hr
Substituting these values into Equation 4-26 for a minimum dt:
= (0.0162)[5094(vTT7)]05=1.2 m
Correct this diameter for a tray spacing of 0.53 m.
From Figure 4-14 read the correction factor as 1.05.
Therefore, the minimum diameter is:
dt= 1.2 (1.05)= 1.26m (4. 13 ft)
Note: that this estimated diameter is a minimum acceptable diameter based on
primary conditions. In practice, a larger diameter based on economic conditions is
usually chosen.
Number of Theoretical Plates or Trays
There are several methods used to determine the number of ideal plates or trays
which are required for a given removal efficiency. These methods, however, can
become quite complicated. One method used is a graphical technique.
The number of ideal plates is obtained by drawing "steps" on an operating
diagram. This procedure is illustrated in Figure 4-15. This method can be rather
time consuming and inaccuracies can result at both ends of the graph.
4-34
-------
(X,
[Note- Lines AB-BC are one theoretical plate
Need a total of 2.3 plates]
X
Figure 4-15. Graphic determination of the number of theoretical plates.
Equation 4-27 is a simplified method to estimate the number of plates. Equation
4-27 can only be used if both the equilibrium and operating lines for the system
are straight. This is a valid assumption for most air pollution control systems.
Equation 4-27, taken from Sherwood and Pigford (1952), is very similar to Equa-
tion 4-24 for computing the N0c of a packed tower.
(Eq. 4-27)
Y2 — mX.
_
I
mG
~T
In
-L-
\mGn
Equation 4-27 is used to predict the number of theoretical plates required to
achieve a given removal efficiency. The operating conditions for a theoretical plate
assume that the gas and liquid stream leaving the plate are in equilibrium with
each other. This ideal condition is never achieved in practice. A larger number of
actual trays are required to compensate for this decreased tray efficiency.
Three types of efficiencies are used to describe absorption efficiency for a plate
tower: an overall efficiency, which is concerned with the entire column; Murphree
efficiency, which is applicable with a single plate; and local efficiency, which per-
tains to a specific location on a plate. A number of methods are available to
predict these plate efficiencies. These methods are complex and values predicted by
two different methods for a given system can vary by as much as 80% (Zenz, 1972).
4-35
-------
The simplest of the tray efficiency concepts, the overall efficiency, is the ratio of
the number of theoretical plates to the number of actual plates. Since overall tray
efficiency is an oversimplification of the process, reliable values are difficult to
obtain. For a rough estimate, overall tray efficiencies for absorbers operating with
low viscosity liquid normally fall in a 65 to 80% range (Zenz, 1972).
Example 4-6
Calculate the number of theoretical plates required for the scrubber in Example
4-5, using the same conditions as in Example 4-4. Estimate the total height of the
column if the trays are spaced at 0.53 m intervals and assume an overall tray effi-
ciency of 70%.
Solution:
From Example 4-5 and the previous examples the following data is obtained:
Slope of equilibrium line, m = 42.7
Y! = 0.03
Y2 = 0.003
X2 = 0.0
Lm = 204 kg mole/min
Gm = 3.5 kg mole/min
The number of theoretical plates from Equation 4-26 is:
rv.-mX./mftA
LYi-roX, \ L. /
—- -
, I" 0.03-0 / 42.7(3.5)\ 42.7(3.5)1
In 1 — +
= L0.003-0\ 204 / 204 J
f 204 1
142.7(5.5)]
In
= 3.94 theoretical plates
Assuming that the overall plate efficiency is 70%, the actual number of plates is:
3.94
Actual plates = —— =5.6 or 6 plates
^ 0.70 r
The height of the tower is given by:
Z = Np X tray spacing + top height
4-36
-------
The top height is the distance over the top plate which allows the gas-vapor mix-
ture to separate. This distance is usually the same as the tray spacing.
Z = 6 plates (0.53 m) + 0.53 m
= 3.71m
Note: that this height is approximately the same as that predicted for the packed
tower in Example 4-4. This seems logical since both the packed and plate towers
are efficient gas absorption devices. However, due to the many assumptions, no
concrete generalization can be made.
Absorption Equipment
Introduction
The primary function of an absorber is to remove gaseous contaminants from an
exhaust air stream. To accomplish this, absorption equipment is designed to max-
imize the mass transfer rate. In absorption, the rate of mass transferred depends
largely on the surface area of the air and liquid stream exposed to each other.
Absorption proceeds at a finite rate. Increasing the time the two streams are in
contact will increase the potential for absorption to occur. Absorbers are designed
to provide the necessary surface area and sufficient contact time between the gas
and liquid streams.
Although the primary objective of an absorber is to remove gaseous con-
taminants, it can also perform other functions. Removal of paniculate matter is a
prime example. When used for gaseous removal the control device is referred to as
an "absorber". When used for paniculate matter removal, the device is termed a
"wet scrubber". These terms are used interchangeably as they describe the same
piece of equipment. The difference lies in the manner in which the equipment is
operated. Absorption is enhanced by slowing the relative velocity between the gas
and liquid streams (increase contact time). Whereas for paniculate matter
removal, the exact opposite is true. Collection efficiency increases by increasing the
relative velocity of the two streams (decrease contact time). The same piece of
equipment can be used for either gas or paniculate matter removal. But it is
extremely difficult to achieve a high removal efficiency for both pollutants unless
the gaseous contaminant is extremely soluble.
The physical and chemical characteristics of the exhaust gas stream play an
important role in both the selection and proper operation of an absorption system.
The solubility of the gaseous contaminant is the first characteristic to evaluate. If
the gaseous contaminant is very soluble, then high removal efficiencies can be
achieved by almost any absorption device. For a relatively insoluble contaminant
only certain systems may be able to achieve the required removal efficiency. In
some cases a chemical reagent may have to be added to the absorbing liquid to
increase the solubility of the contaminant. These reagents may increase the physical
solubility of the contaminant (i.e. sodium citrate added to absorb SO2) or can
4-37
-------
chemically react with the contaminant (i.e. lime scrubbing of SO2). If a precipitate
is formed by a chemical reaction when a reagent is used, plugging or corrosion
problems may arise.
The temperature of the exhaust stream is another important characteristic which
affects absorption. The solubility of a gas decreases with an increase in operating
temperature. As the temperature increases, so does the kinetic energy of the gas
molecules in solution. At these higher states of energy the gas molecules will come
out of solution. This loss of solubility at higher temperatures necessitates that some
gas streams be cooled before effective absorption occurs. For example, in certain
FGD systems the exhaust gases from the boiler must be cooled from 150°C (300 °F)
to approximately 50 °C (125°F) to achieve the desired SO2 removal efficiency. This
is accomplished by adding inlet sprays or adjusting the liquid flow rate to the
absorber. The temperature of the exhaust stream also affects the size of the absorp-
tion system. Decreasing the temperature decreases the volume of gases which must
be handled. This decreases the size of the absorption system.
Selection of the proper absorbing liquid is based on the efficiency required and
the liquid cost. Water is the usual choice because many gaseous contaminants are
soluble in it, it is readily available, and relatively inexpensive. The following
properties must also be kept in mind when selecting a liquid:
Gas solubility: high solubility increases the absorption rate and minimizes the
quantity of liquid needed.
Volatility: low volatility of the liquid will reduce the amount of vapor that is
lost in the existing gas stream.
Viscosity: low viscosity promotes rapid absorption rates, improves flooding
characteristics, and lowers the pressure drop.
Chemical stability: the absorbent should not degrade but remain effective
throughout its useful lifetime.
Flammability: if at all possible, the liquid should be nonflammable, non-
corrosive, nontoxic, and inexpensive.
An important factor in the successful operation of any absorber is to initially
select the proper construction materials. Quenching hot gases to their saturation
temperatures forms corrosive acids. Depending on the substances present in the
exhaust stream, sulfuric, hydrochloric, or hydrofluoric acids may be formed. The
presence of these acids can cause severe corrosion problems unless special materials
of construction are used. The construction materials typically used include rubber
and PVC lined absorber vessels. Currently there is a trend toward using more
absorbers made of reinforced plastics when the gaseous contaminant and/or the
scrubbing solution is corrosive (McDonald, 1977).
Spray Towers or Chambers
Spray towers, the simplest devices used for gas absorption, consist of an empty
tower and a set of nozzles to spray liquid. A spray tower is similar in operation to
spraying water in an open barrel. Typically, the contaminant gas stream enters the
bottom of the tower and passes up through the device while liquid is being sprayed
4-38
-------
at one or more levels by nozzles. The flow of liquid and gas streams in opposite
directions is referred to as counter current flow. Figure 4-16 illustrates the operation
of a typical countercurrent flow spray tower.
Clean gas out
Spray nozzles
Dirty gas in
Figure 4-16. Spray tower.
To provide a large liquid surface for contacting the gas, nozzles are arranged to
wet the entire cross section of the tower with fine liquid droplets. Theoretically, the
height between levels of nozzles and/or the bottom of the tower determines the
residence time. In practice, physical laws limit the removal efficiency of spray
towers. After falling short distances the liquid droplets tend to agglomerate or hit
the sides of the tower. Both of these effects reduce the total liquid surface in con-
tact with the gas stream and the residence time. Therefore, spray chambers are
limited to applications where the gases are extremely soluble or a high removal effi-
ciency is not required.
The main advantage of spray chambers is that they are completely open; they
have no internals except for the spray nozzles. Therefore, they have a very low
pressure drop, about 1.25 to 4 cm (0.5 to 1.5 in.) of water over the tower. If an
entrainment separator is used it will add another 2.5 to 5 cm (1 to 2 in.) of water
to the total pressure drop. An entrainment separator is used to prevent mist from
exhausting out the stack. Spray towers range in size from 42 to 170,000 mVhr (25
to 100,000 cfm) with 34,500 mVhr (20,000 cfm) as a typical size (EPA, 1972).
4-39
-------
Spray towers are primarily used for gas conditioning—cooling, humidifying, or for
the first stage paniculate matter and gas removal. But they can be used effectively
for gas absorption if the contaminant to be removed is highly soluble. For example,
spray towers are used to remove HC1 gas from the tail gas exhaust in the manufac-
turing of hydrochloric acid. In the production of superphosphate used in manufac-
turing fertilizer, SiF4 and HF gases are vented from various points in the processes.
Spray towers have been used to remove these highly soluble compounds. Spray
towers are also used for odor removal in bone meal and tallow manufacturing
industries by scrubbing the exhaust gases with a solution of KMnO4. Due to their
ability to handle large volumes of exhaust gases in corrosive atmospheres, spray
towers are also used in a number of flue gas desulfurization systems as first or
second stage removal devices.
No universally acceptable mathematical equation or correlation is presently
available for sizing spray towers. Although they are simple devices, variables such
as liquid drop size, settling velocity and residence time vary considerably with
height or location in the tower. Tower designs are based on experimental or opera-
tional data from similar systems. However, certain limiting factors must be
balanced against each other. First, the gas velocity through the tower must be kept
low, around 0.3 to 1.0 m/sec (1 to 3 ft/sec) or excessive liquid entrainment occurs.
This implies using a large diameter tower to keep the gas velocity low. But if the
diameter is large, compared to the distance between sprays, back mixing occurs.
Back mixing is a condition where the gas changes direction or swirls within the
tower and removal efficiency is decreased.
The smaller the liquid droplet size, the greater the rate of absorption, since this
increases the surface area available for absorption. High pressure spray nozzles are
used to produce fine liquid droplets. These spray nozzles consume more power in
pumping liquid than a packed tower or venturi scrubber for the same liquid rate.
The fine openings of the spray nozzles are subject to erosion and plugging
problems, especially if scrubbing liquid is recirculated.
Another parameter which characterizes spray tower operation is the liquid to gas
ratio (L/G). By increasing the L/G ratio, absorption efficiency can be directly
increased. Figure 4-17 illustrates how the percent SO2 removal is affected by the
L/G ratio. Figure 4-17 was developed by Ontario Hydro Electric from data for SO2
removal in a spray tower using limestone on a 4000 cfrn, 30 MW test unit. The
disadvantage of increasing the liquid flow rate is that is adds to the operating costs
both in liquid usage and power consumed.
4-40
-------
100
75
Si
o
I
50
25
V,= 8.6 ft/sec
P = 15 psig
25 50 75
L/G, gal/1000 ft3 (120°F)
100
Figure 4-17. Spray tower efficiency vs. L/G.
Where there is limited space in a plant, spray chambers can operate by cross-
current or cocurrent flow. In crosscurrent absorbers the gas flow is perpendicular
to the liquid flow. In cocurrent absorbers the gas and liquid flow in the same
direction. Because the gas stream does not "push" against the liquid stream as in
countercurrent flow, much higher gas stream velocities can be used. Higher gas
stream velocities mean that the size of the unit can be reduced. However, cross-
current or cocurrent spray towers are not as efficient as the countercurrent type.
Countercurrent flow exposes the most dilute portion of the contaminant gas stream
to the freshest liquid throughout the entire tower length.
Another spray device sometimes used for gas absorption is the cyclonic scrubber
pictured in Figure 4-18. The gas stream enters the device tangentially which causes
it to swirl up through the tower in a corkscrew motion. Liquid is sprayed from the
center outward into the swirling gas stream (vortex). This configuration allows for
the use of higher gas velocities without excessive liquid entrainment. The liquid
droplets are forced to the walls of the device by centrifugal force before they can
be carried out the top of the device. Due to the short residence time of the liquid
droplets in the gas stream, removal efficiency is very limited.
4-41
-------
Figure 4-18. Cyclonic spray scrubber.
High Energy Absorbers
High energy scrubbers (Venturis and ejectors) are primarily used for high efficiency
particulate matter collection in the submicrometer range. These devices can also be
used for absorption although they have a limited gas collection efficiency due to
the short residence time of gas-liquid contact. For the collection of particulate
matter with high energy scrubbers, increased energy input increases collection effi-
ciency. For gas absorption, almost the direct opposite holds true. An increase in
pressure drop decreases residence time, thus reducing the amount of mass transfer.
The most common high energy scrubber is the venturi. A typical venturi con-
figuration is shown in Figure 4-19. A venturi consists of a converging section
through which the gas enters and is accelerated to a high velocity. The gas shears
the liquid from the walls of the throat section producing an enormous number of
fine droplets. The throat section is a constant cross sectional area where the
majority of absorption occurs. The gas then exits through an expander section
where it is slowed down, reducing energy consumption. Simple Venturis, as shown
in Figure 4-19, can be of two types depending on where the liquid is injected. If
the liquid is injected at the throat section it is referred to as a nonwetted approach
venturi. If the liquid is introduced through pipes at the entrance of the converging
4-42
-------
section it is called a wetted approach venturi. Wetted approach Venturis are used
to saturate the incoming gas stream. This eliminates the possibility of having a wet-
dry interface area at the entrance to the throat. At these wet-dry interface areas,
scale build-up can eventually cause plugging problems. The wetted approach is
usually more expensive than nonwetted approach units, so if the inlet gas is already
saturated a nonwetted approach may be preferred (Brady, 1977).
Converging section
Diverging section
Liquid inlet
Throat
Figure 4-19. Typical venturi configuration.
4-43
-------
The simple venturi shown in Figure 4-19 is adequate for flows up to 88,000
mVhr (40,000 acfm). At flows greater than this it is difficult to achieve uniform
liquid distribution unless additional weirs or baffles are added. To handle the
larger gas flows, manufacturers have gone to a rectangular configuration with a
long narrow throat as shown in Figure 4-20 (Brady, 1977). Other types of Venturis
can have variable throat openings. An adjustable plumb bob in the throat sec-
tion is used to accommodate variable flows (Figure 4-21). Water is sprayed on the
plumb bob in the throat section, spreads out and is sheared off by the gas. The
plumb bob moves up or down depending on the gas flow rate and the desired
pressure drop.
Liquid sprays
Cyclonic separator
Flooded elbow
Figure 4-20. Spray venturi.
4-44
-------
Liquid
injection
nozzle
Plunger
(plumb bob)
Figure 4-21. Adjustable throat venturi.
Venturi scrubbers normally operate with high gas side pressure drop, ranging
from 50 to 250 cm (20 to 100 in.) of water. These high pressure drops result both
in high energy costs and high maintenance costs for the fans needed to move the
gas stream. The inlet gas velocity is greatest through the throat section of the ven-
turi. In the throat, gas velocities can vary between 30 to 125 mVs (100 to 400
ft/sec). These high velocities are good for particulate matter removal, but result in
insufficient contact time for most absorption applications. High gas velocities also
make it mandatory that some type of entrainment separator be used (Figure 4-20).
The entrainment separator will eliminate any liquid carryout in the exhaust from
the scrubbing system.
To overcome the short residence time, most venturi systems operate at high
liquid-to-gas (L/G) ratios when used for gaseous emission control. Typical liquid-to-
gas ratios for Venturis used as absorbers range between 47 and 105 liters per
min/100 m3 per hr (20 to 40 gpm/1000 cfm) and in some instances up to 195 liters
per min/1000 m3 per hr (80 gpm/1000 cfm). For particulate matter removal, Ven-
turis normally operate between 4.7 and 38 liters per min/1000 m3 per hr (2 to 15
gpm/1000 cfm).
4-45
-------
Another type of venturi is the jet or ejector venturi (Figure 4-22). This device
uses high pressure spray nozzles (up to 100 psig) aimed at the throat section of a
venturi. Using high pressure sprays in this manner serves two purposes. First the
movement of the liquid creates a suction (vacuum) which pulls the gas stream
through the venturi. This eliminates the need of a fan or blower to move the gas
stream. Secondly, the high pressure sprays, along with the venturi effect, form
numerous fine liquid droplets which provide a high degree of turbulence between
gas and liquid phases. This limits contact time, therefore, absorption efficiency is
usually low.
Liquid
injection
nozzle
P
Figure 4-22. Jet or ejector venturi.
Ejector Venturis operate at high liquid-to-gas ratios (around 100 gpm/1000 cfm).
The gas side pressure drop usually ranges between 10 to 20 cm (4 to 8 in.) of
water. The overall power consumption, however, is much higher, due to liquid
pumping requirements. Ejector Venturis can be designed for smaller gas flow rates
(under 1000 cfm) than most other absorbers are capable of handling.
Venturi scrubbers are the most widely used wet collector for paniculate matter
removal. Unless a gaseous contaminant is extremely soluble, Venturis are seldom
used strictly as an absorber. A variety of Venturis are used by power plants to
remove both flyash and sulfur dioxide. To achieve the required SO2 removal effi-
ciencies Venturis are usually used in combination with another absorber. Venturis
have been used with spray towers, packed towers, and other Venturis to obtain the
4-46
-------
desired removal efficiency. Venturi scrubbers are also used to control emissions
from the smelting of lead, zinc and copper. Smelting operations typically produce
emissions of very fine acid mists, high concentrations of sulfur dioxide, and a
variety of particulate matter.
Packed Column
The packed column (tower) is the most common scrubber used for gas absorption.
Packed columns disperse the scrubbing liquid over packing material which provides
a large surface area for continuous gas-liquid contact. Packed towers are classified
according to the relative direction of gas-to-liquid flow.
The most common packed tower is the countercurrent (gas-to-liquid) flow tower,
shown in Figure 4-23. The gas stream being treated enters the bottom of the tower
and flows upward over the packing material. Liquid is introduced at the top of the
packing by sprays or weirs and flows downward over the packing material. This
flow arrangement results in the highest theoretical achievable efficiency. The most
dilute gas is contacted with the purest absorbing liquor, providing a constant,
maximized concentration difference (driving force) for the entire length of the
column.
Mist eliminator
Liquid sprays
Liquid outlet
Figure 4-23. Countercurrent packed tower.
4-47
-------
Pressure drop in the gas phase ranges from 0.63 to 2.54 cm (0.25 to 1 inch per
ft) of water per meter of column packing at 40 to 70% of the flooding velocity.
This absorber cannot handle large variations in liquid or gas flow rates or loading
and flooding will occur. Loading and flooding conditions were described in the
preceding design section on packed tower diameter.
In another flow arrangement used with packed towers, cocurrent flow, both the
gas and liquid phases enter at the top of the absorber and move downward over
the packing material. This allows the absorber to be operated at higher liquid and
gas flow rates since flooding is not a problem. The pressure drop is lower than with
countercurrent flow since both streams move with gravity. The major disadvantage
is that removal efficiency is very limited due to the decreasing driving force (con-
centration differential) as the streams travel down through the column. This limits
the areas of application for cocurrent absorbers. They are used almost exclusively
in situations where limited equipment space is available since the tower diameter is
smaller than a countercurrent or plate tower for equivalent flow rates.
In a cross flow absorber, the gas stream flows horizontally through the packed
bed which is irrigated by the scrubbing liquid flowing down through the packing
material perpendicular to the gas flow. A typical cross flow absorber is shown in
Figure 4-24. Inlet sprays aimed at the face of the bed may also be included. These
sprays scrub both the entering gas and the face of the packed bed. The leading
face of the packed bed is slanted in the direction of the on-coming gas stream as
shown in Figure 4-24. This insures complete wetting of the packing by allowing the
liquid at the front face of the packing time to drop to the bottom before being
pushed back by the entering gas.
Liquid sprays
Packing
Figure 4-24. Cross flow packed tower.
4-48
-------
Cross flow absorbers are smaller and have a lower pressure drop than any other
packed or plate tower for the same application (removal efficiency and flow rates).
In addition, they are better suited to handle exhaust streams with heavy paniculate
matter concentration. By adjusting the liquid flow rate, incoming paniculate
matter can be removed and washed away in the front half of the bed. This also
results in a liquid savings by enabling the absorber to use less liquid in the rear
sprays. This practice is carried one step further by actually constructing the
absorber into sections as shown in Figure 4-25. The front section can be equipped
with water sprays and used for paniculate matter removal. In the second section,
sprays may contain a reagent in the scrubbing liquor for gas removal. The last sec-
tion is left dry to act as an entrainment separator. Cross flow absorbers do require
complex design procedures since concentration gradients exist in two directions in
the liquid; from top to bottom and from front to rear.
Liquid sprays
n\
Figure 4-25. Three bed cross flow packed tower.
Packed towers are most suited to applications where a high gas removal effi-
ciency is required and the exhaust stream is relatively free from paniculate matter.
In the production of both sulfuric and hydrochloric acids, packed towers are used
to control tail and exhaust emissions (SO2 and HC1 respectively). The scrubbing
liquor for these processes can be a weak acid solution, with the spent liquor from
the packed tower sent back to the process. Packed towers are also used to control
HC1 and H2SO4 fume emissions from pickling operations in the primary metals
industry. They are also used to control odors in rendering plants, petroleum
refineries, and in sewage treatment plants. For odor control, the scrubbing liquor
usually contains an oxidizing reagent such as potassium permanganate or sodium
hypochlorite. In these applications, an acid backwash must be added if a
precipitate is formed or plugging can be a problem. The gas flow rate through
packed towers can vary from 17 to 51,000 mVhr (10 to 30,000 cfrn) with 8,600
mVhr (5100 cfm) as an average size (EPA, 1972).
4-49
-------
Packing Material
Packing material is the heart of the absorber. It provides the surface over which
the scrubbing liquid flows, presenting a large area for mass transfer to occur.
Packing material represents the largest material cost of the packed tower. Pictured
in Figure 4-26 are some of the more commonly used packings which are made in
numerous geometric shapes and sizes. These materials were originally made of
stoneware, porcelain, or metal; but presently a large majority are being made of
high-density thermoplastics (polythylene and polypropylene). A specific packing is
described by its trade name and overall size. For example, a column can be packed
with 5 cm (2 inch) Raschig rings or 2.5 cm (1 inch) Tellerettes. The overall dimen-
sion of packing materials normally range from 0.6 to 10 cm (V4 to 4 inches).
Raschig Ring
Pall ring
Berl saddle
Intalox saddle
Tellerette
Figure 4-26. Common packing materials.
The specific packing that is selected depends on the nature of the contaminants,
geometric mode of contact, size of the absorber, and scrubbing objectives. The
following factors provide a general guide for selecting packing materials
(McDonald, 1977). (For some specific packing materials, these factors are listed in
Table 4-3).
Cost: Generally plastic packings are cheaper than metal with ceramic being the
most expensive. Cost of packings are given in dollars per cubic meter ($/ms).
Low pressure drop: Pressure drop is a function of the volume of void space in a
tower when filled with packing. Generally, the larger the packing size, the smaller
the pressure drop.
Corrosion resistant: Ceramic or porcelain packing is commonly used in a very
corrosive atmosphere.
Large specific area: A large surface area per cubic foot of packing, mVm3
(ftVft3), is desirable for mass transfer.
4-50
-------
Structural strength: Packing must be strong enough to withstand normal loads
during installation, service, physical handling, and thermal fluctuations. Ceramic
packing is subject to cracking under sudden temperature changes.
Weight: Heavier packing may require additional support materials or heavier
tower construction. Plastics have a big advantage in this area since they are much
lighter than either ceramic or metal packings.
Design flexibility: The efficiency of a scrubber changes as the liquid and gas
flow rates are varied. Packing material must be able to handle the process changes
without substantially affecting the removal efficiency.
Packing material may be arranged in an absorber in either of two ways. The
packing may be dumped into the column randomly or in certain cases system-
atically stacked, as bricks are laid atop each other. Randomly packed towers pro-
vide a higher surface area, mz/m3 (ft2/ft3), but also cause a higher pressure drop
than stacked packing. In addition to the lower pressure drop, the stacked packing
provides better liquid distribution over the entire surface of the packing. The large
installation costs required to stack the packing material usually make it impractical
unless high flow rates are required.
Liquid Distribution
As stated previously, one of the keys to effective absorber operation is to intimately
contact the gas stream with the liquid stream. This contact must be maintained
throughout the entire column length. No packing material will adequately
distribute liquid which is poured onto it at only one point. Liquid introduced into
the absorber in this manner tends to flow down over a relatively small cross section
of the tower diameter. Known as channeling, the liquid flows in little streams down
through the tower without wetting the entire packing area. Liquid should be
distributed over the entire cross sectional top of the packing. This is commonly
achieved by weirs or feed tube arrangements as shown in Figure 4-27. Although
liquid sprays are also used, weirs or tube arrangements require less power, are
more flexible to changing liquid loads, and handle recycled liquid better than
sprays.
Once the liquid is distributed over the packing, it flows down, by the force of
gravity, through the packing, following the path of least resistance. The liquid
tends to flow towards the tower wall where the void spaces are greater than in the
center. Once the liquid hits the wall, it flows straight down the tower from that
point (channels). A way must be provided to redirect the liquid from the tower
wall back to the center of the column. This is usually done by using liquid
redistributors which funnel the liquid back over the entire surface of packing. It is
recommended that redistributors be placed at intervals of no more than 3 meters
(10 feet) or 5 tower diameters, whichever is smaller (Zenz, 1972).
Uniform distribution of the feed gas is also very important. This is accomplished
by properly designing the support trays which hold up the packing material. If the
tower is broken into sections, each support grid would act as a distribution baffle.
4-51
-------
(a) riser orifice
(b) notched riser
(c) trough and weir type (d) perforated pipe
Figure 4-27. Liquid feed distributors.
Plate Tower
A plate tower is a vertical column with one or more plates (or trays) mounted
horizontally inside. The gas stream enters at the bottom and flows upward passing
through openings in the plates. Liquid enters at the top of the tower and travels
across each plate, then through a downcomer to the plate below until it reaches the
bottom of the tower. Mixing occurs as the gas bubbles up through the layer of
liquid covering each plate. Figure 4-28 illustrates the flow pattern of a typical plate
tower.
4-52
-------
Liquid
downcomer
Mist eliminator
Liquid inlet
Plate or tray
Figure 4-28. Plate tower.
Mass transfer occurs as the gas stream bubbles up through the openings and into
the liquid layer on the plate. The smaller the bubbles and the more numerous they
are, the more effective the mass transfer. The formation of smaller and more
numerous bubbles promotes better mixing and exposes a larger gas-to-liquid sur-
face area. As the gas disengages from the liquid, a froth is formed atop the liquid.
Ideally, the gas and liquid leaving each plate are in equilibrium with each other at
the conditions of that particular plate. Each plate acts as a separate absorption
stage. The amount of absorption that can occur at one plate or stage is referred to
as a transfer unit.
4-53
-------
The function of the plates or trays is to disperse the gas into numerous bubbles,
exposing a large surface for mass transfer. As with packing material a variety of
different plates are in use, most commonly:
Bubble cap: Vapor enters the cap through risers and out of the slots as bubbles
into the surrounding liquid on the plate (Figure 4-29a). The main advantage is
that this design can handle a wide range of gas and liquid rates without affecting
efficiency. Plugging and corrosion can be a problem.
Sieve trays: Vapor rises through small holes 0.60 to 2.50 cm in diameter (V4 to 1
inch) in the plate and bubbles through the liquid on the plate (Figure 4-29b).
Trays can contain 600 to 3000 holes per square foot of surface. The installed costs
are lower than for any other type of plate and the fouling tendency (with large
holes) is lower.
Valve trays: Vapor flows up through small holes and lifts up metal valves or
caps that cover the openings in the plate. The valves are restrained by legs which
limit vertical movement (Figure 4-29c). The liftable caps act as variable orifices;
adjusting the opening proportional to vapor flow. Caps are arranged in different
weights to provide flexibility. Valve trays are more expensive than sieve trays but
cheaper than bubble cap trays. They can handle high liquid-to-gas ratios. Wear or
corrosion of retaining legs can be a problem.
(a) bubble cap plate
(b) sieve tray
(c) float valve
Figure 4-29. Scrubbing plates or trays.
Plate towers, like packed towers, can achieve a high removal efficiency even if
the gaseous contaminant is relatively insoluble. Therefore, plate towers are used to
control emissions from many of the same processes that use packed towers, such as
from acid plants, fertilizer production plants, chemical industry and petroleum
refineries.
4-54
-------
Several advantages and disadvantages must be considered when choosing a plate
tower instead of a packed tower for a given control operation. The following list
gives some factors used in comparing plate towers to packed towers.
1. Plate towers are able to handle particulate matter and other solids better
than packed towers. Maintenance openings can be installed so that the plates
may be easily cleaned.
2. Plate towers are chosen for operations that require a large number of transfer
units or that must handle large gas volumes. Packed towers can experience
channeling problems if the diameter or height of the tower is too large.
Redistribution trays must be installed in large diameter and tall packed
towers to avoid channeling.
3. The total weight of a plate tower is less than that of a comparable packed
tower.
4. Packed towers are much cheaper to construct if corrosive substances are to be
handled. Packed towers can be constructed with a fiberglass reinforced
polyester shell which is generally about half the cost of a carbon steel plate
tower.
5. Plate towers can handle volume and temperature fluctuations better than
packed towers. Expansion or contraction due to temperature changes can
crush or melt packing material.
Mobile Packed Beds
Mobile packed beds are very similar to packed towers. However, instead of being
stationary, they provide a zone of moving packing material. The gas and the liquid
streams are mixed in this moving zone where mass transfer occurs. These scrubbers
are primarily used when both particulate and gaseous contaminant removal is
necessary. They provide the mass transfer efficiency of a packed or plate tower
without the plugging problems. They do not, however, have the efficiency of a
packed tower per energy unit consumed. Mobile packed beds are classified as either
a flooded bed or a fluidized bed, depending on the degree of packing movement.
The gas and liquid streams can flow either cocurrently or countercurrently.
4-55
-------
The flooded bed absorber pictured in Figure 4-30 is made up of one section of
mobile spheres 10 to 20 cm (4 to 8 inches) deep. The spheres are usually made of
plastic, however, glass or marble spheres are also used. The gas stream enters from
the bottom while liquid is sprayed from the top and/or bottom over the packing.
The gas velocity is such that it causes the packing to rotate in a gentle rubbing
motion, but does not fluidize the packing. The rotating movement acts as a self-
cleaning mechanism plus enhances gas and liquid mixing. Bubbles formed in the
bed create a layer of froth over each section of bed which is about twice as high as
the bed itself.
Mist eliminator
Packing
Liquid sprays
Figure 4-30. Flooded bed absorber.
A fluidized bed absorber is very similar to a flooded bed absorber. The dif-
ference is in the degree of movement of the packing. In a fluidized bed absorber,
the gas velocity is such (1.8 to 2.4 m/sec) (6 to 8 ft/sec) that it keeps the packing
in constant motion between a lower and upper retaining grid. This is shown in
Figure 4-31. The packing is made of either polypropylene or polyethylene plastic
balls that are hollow, resembling a ping pong ball. The packed sections are usually
0.3 to 0.6 m (1 to 2 ft) thick with a froth zone about 0.6 meters (2 ft) thick above
the packing.
4-56
-------
" U Q^ 6 ,-[
~L °oc*0 o0 no o °
°o o So 00rirt0 o
o o o
Mist eliminator
Liquid sprays
Fluidized packing
Figure 4-31. Fluidized bed absorber.
The pressure drop in these mobile packed beds ranges from 5 to 15 cm (2 to 6
inches) of water per stage of packing. The capacity of these units can be 5 to 6
times that of a packed or plate tower of the same diameter due to the high gas
velocities through the units (Bethea, 1978). Mobile packed beds are used to control
emissions from exhaust streams that require a high degree of gaseous contaminant
removal and paniculate matter removal. Mobile packed beds have been used to
control emissions from Kraft pulp mills, cupola furnaces, and aluminum foundries.
These units have also been used for SO2 and flyash control from large power
plants.
Entrainment Separators
Gas moving at high velocities that mixes with a liquid will entrain drops of that
liquid. The liquid droplets must be removed from the gas stream before being
exhausted to the atmosphere. Entrainment separators are used to remove the liquid
droplets. Although the major function of an entrainment separator is to prevent
liquid carryover, it also performs additional scrubbing and recovers the scrubbing
liquor which saves on operating costs. Entrainment separators are therefore usually
an integral part of any wet scrubbing system.
4-57
-------
Entrained liquid droplets vary in size depending on how the droplets were
formed. Drops that are torn from the body of a liquid are large (10 to 100
micrometers in diameter); whereas drops that are formed by a chemical reaction or
by condensation are on the order of 5 micrometers or less in diameter. Numerous
types of entrainment separators are capable of removing these droplets. Those most
commonly used for air pollution control purposes are the cyclone, wire mesh pads,
and blade types.
The cyclone (centrifugal) separator is a cylindrical tank with a tangential inlet or
turning vanes. The inlet or vanes impart a swirling motion to the droplet laden gas
stream. The droplets are thrown outward by centrifugal force to the walls of the
cylinder. Here they coalesce and drop down the walls to a central location and are
recycled to the absorber. These units are simple in construction, having no moving
parts. Therefore, they have few plugging problems. Good separation down to the 10
micrometer range can be expected. The pressure drop across the cyclone is 10 to
15 cm (4 to 6 inches) of water for a 98% removal efficiency of droplets down to the
20 to 25 micrometer range. Cyclonic separators are commonly used with venturi
scrubbers as seen in Figures 4-20 and 4-21.
Wire mesh or plastic pad separators consist of woven material about 10 to 15 cm
(4 to 6 inches) thick that fit across the entire diameter of the scrubber (Figure
4-32). The mesh allows droplets to impact on the material surface, agglomerate
with other droplets and drain off by gravity. The pad is usually slanted (no more
than a few degrees) to permit the liquid to drain off. Essentially 100% collection of
droplets larger than 3 micrometers is obtained with pressure drops of approxi-
mately 10 to 15 cm (4 to 6 inches) of water (the pressure drop is dependent on
depth and compaction of fibers). The disadvantage with mesh pads is that their
small passages are subject to plugging.
Chevron
Figure 4-32. Mist eliminators.
4-58
-------
Impingement blade separators employ blades bent into "S" curves or a chevron
configuration (Figure 4-32). As gas passes between the blades, it is forced to travel
in a zigzag pattern. The liquid droplets cannot follow the flow lines and impinge
on the blade surface, coalesce and fall back into the scrubber chamber (illustrated
in Figure 4-30). The pressure drop is approximately 6.4 cm (2.5 inches) of water
for the capture of droplets as small as 5 micrometers in diameter.
References
Bethea, R. M., 1978. Air Pollution Control Technology. New York: Van Nostrand
Reinhold.
Bhatia, M. V. 1977. Packed Tower and Absorption Design. In Air Pollution Con-
trol and Design Handbook. P. N. Cheremisinoff and R. A. Young, eds. New
York: Marcel Dekker, Inc.
Brady, J. D. and L. K. Legatski. 1977. Venturi Scrubbers. In Air Pollution Con-
trol and Design Handbook. P. N. Cheremisinoff and R. A. Young, eds. New
York: Marcel Dekker, Inc.
Calvert, S., 1977. Scrubbing. In Air Pollution. Vol. IV. Engineering Control of
Air Pollution, A. C. Stern, ed. New York: Academic Press.
Environmental Protection Agency (EPA). 1973. Air Pollution Engineering Manual,
AP40. Research Triangle Park, NC.
Environmental Protection Agency (EPA). 1972. Wet Scrubber System Study.
NTIS Report PB-213016. Research Triangle Park, NC.
Marchello, J. M. 1976. Control of Air Pollution Sources. New York: Marcel
Dekker, Inc.
McCabe, W. L. and Smith, C. J. 1967. Unit Operations of Chemical Engineering.
New York: McGraw Hill Book Co.
McDonald, J. W. 1977. Packed Wet Scrubbers. In Air Pollution Control and
Design Handbook. P. N. Cheremisinoff and R. A. Young, eds. New York:
Marcel Dekker, Inc.
Perry, J. H. ed. 1973. Chemical Engineers Handbook, 5th ed. New York:
McGraw Hill Book Co.
Sherwood, T. K. and Pigford, R. L. 1952. Absorption and Extraction. New
York: McGraw Hill Book Co.
Theodore, L. and Buonicore, A. J. 1975. Industrial Control Equipment for
Gaseous Pollutants. Vol. I. Cleveland: CRC Press.
4-59
-------
Treybal, R. E. 1968. Mass Transfer Operations, 2nd ed. New York: McGraw
Hill Book Co.
Whitman, W. G. 1923. Chem. and Met Eng. 29:147).
Zenz, F. A. 1972. Designing Gas Absorption Towers. Chem. Eng. 79:120-138
(November).
4-60
-------
Chapter 5
Adsorption
Introduction
During adsorption, one or more gaseous components are removed from an effluent
gas stream by adhering to the surface of a solid. The gas molecules being removed
are referred to as the adsorbate, while the solid doing the adsorbing is called the
adsorbent. Adsorbents are highly porous particles. Adsorption occurs on the inter-
nal surfaces of the particles as illustrated in Figure 5-1.
Adsorbent
Adsorbate
Figure 5-1. Vapor adsorbed into pores of activated carbon.
The attractive forces which hold the gas to the surface of the solid are the same
that cause vapors to condense (van der Waals' forces). All gas-solid interfaces
exhibit this attraction, some more than others. Adsorption systems use materials
which are highly attracted to each other to separate these gases from the non-
adsorbing components of an air stream. For air pollution control purposes, adsorp-
tion is not a final control process. The contaminant gas is merely stored on the sur-
face of the adsorbent. After it becomes saturated with adsorbate, the adsorbent
must either be disposed of and replaced, or the vapors must be desorbed. Desorbed
vapors are highly concentrated and may be recovered more easily and more
economically than before the adsorption step.
Traditionally, adsorption has been used for air purification and solvent
recovery. Air purification processes are those in which the contaminant is present
in trace quantities (less than 1.0 ppm) but is highly odorous or toxic. Systems used
5-1
-------
for air purification are small thin bed adsorbers. When the bed becomes saturated
with contaminant, it is taken out and replaced. Solvent recovery processes require
much larger systems and are usually designed to control organic emissions whose
concentrations are greater than 1000 ppm. This has been the point where the
recovery value of the solvent could justify the expense of the large adsorption-
desorption system. Currently, adsorption is used as a method of recovering valuable
organic vapors from flue gases at all concentration levels. This is due to present
regulations limiting volatile organic emissions and the higher costs of solvents.
Theory of Adsorption
Mechanism of Adsorption
Adsorption occurs by a series of three steps. In the first step, the contaminant dif-
fuses from the major body of the air stream to the external surface of the adsor-
bent particle. In the second step, the contaminant molecule migrates from the
relatively small area of the external surface (a few mVg) to the pores within each
adsorbent particle. The bulk of adsorption occurs in these pores because the
majority of available surface area is there (hundreds of mVg). In the third step,
the contaminant molecule adheres to the surface in the pore. Figure 5-2 illustrates
this overall diffusion and adsorption process.
Step 1: diffusion to
adsorbent surface
. .
••
Step 2: migration into
pores of adsorbent
Contaminant molecules
1
.•:V .*•/-/.-*-:•>:
•••.*•. • • • • • • • • •
Step 3: monolayer
buildup of adsorbate
•* ••* • •
• • . ^.. •
Figure 5-2. Mechanism of adsorption.
5-2
-------
The purpose of analyzing the mechanism of adsorption is to determine which
step controls the overall process. By describing this step mathematically, adsorber
performance can be predicted from physical data. The actual adsorption of a
molecule, step 3, proceeds relatively quickly compared to steps 1 or 2. Therefore,
step 3 can be ignored when developing design equations. Steps 1 and 2 are both
diffusional processes. They involve the transport of the adsorbate through a carrier
gas phase to an adsorption site. In the first step, diffusion occurs because of a con-
centration difference. The rate of mass transferred by this type of diffusion can be
predicted from Equation 5-1.
Equation 5-1 is based on a film resistance theory of bulk diffusion as presented
in Chapter 4 on absorption. Bulk diffusion assumes that the only resistance a gas
molecule encounters in movement through the carrier gas stream occurs during col-
lisions with other gas molecules.
(Eq.5-1) N,~ W(p-p.)
QB
Where: NA = rate of mass transfer, kg mol/s
kg = local mass transfer coefficient, kg mol/s»m2»Pa
/3 = void area between adsorbent granules, mVm3
A = surface area of adsorbent, mVkg
p = partial pressure of adsorbate in gas phase, Pa
p, = partial pressure of adsorbate at the gas-solid interface, Pa
g ,B = bulk density of adsorbent, kg/m3
The mass transfer coefficient (kf) is a function of the velocity, viscosity, and density
of the carrier gas stream; the diffusivity of the gas molecule that is being adsorbed;
and the diameter of the adsorbent. Equations to estimate the transfer coefficient
based on these parameters can be found in Perry (1973).
Once the gas molecule has reached the external surface of the adsorbent, it must
diffuse (move) into the pores. Diffusion in the pores of the adsorbent can occur by
a number of different diffusion mechanisms depending on the size of the pore.
When the pores are large, bulk diffusion predominates. As the pores begin to nar-
row, collisions with the wall of the adsorbate become more likely than inter-
molecular collisions. The gas molecules strike the wall, remain for a short time,
then return to the gas phase. This is termed Knudson (or molecular) diffusion and
occurs much more slowly than does bulk diffusion for a given pore length. Finally,
in the smallest pores, surface diffusion is the predominant mechanism of gas
transport. Gas molecules can either migrate along the surface of the solid or jump
between adsorption sites.
Due to these varied mechanisms by which diffusion occurs, mass transfer rates in
the pores are extremely difficult to predict. Unless extensive data are available con-
cerning the specific adsorption application, determining the rate-controlling step
(step 1 or step 2) is impossible.
5-3
-------
One approach to determining the mass transfer rate is to rewrite Equation 5-1 in
terms of an overall mass transfer coefficient.
(Eq. 5-2) Nx = Koca(p-p*)
Where: a = external adsorbent area, mVm3
p* = partial pressure in equilibrium with the surface concentration
of adsorbate, Pa
p = partial pressure of adsorbate in the gas phase, Pa
Koc = overall mass transfer coefficient, kg mol/h«m2»Pa
The overall mass transfer coefficient represents the resistance to molecular motion
encountered both outside and inside the pore.
(Eq. 5-3) — L- = JL + _L
kg k.
Where: k, = local mass transfer coefficient for combined surface migration and
pore diffusion, kg mol/h»mz»Pa
The local mass transfer coefficient cannot be satisfactorily predicted from basic
theory at the present time. However, it (and therefore Koc) can be determined with
some certainty from experimental data. Therefore, Equation 5-2 still does not give
a simple and accurate means of predicting adsorber performance from physical
data. What Equation 5-2 does show is that the equilibrium partial pressure of the
adsorbate (p*) determines the theoretical minimum adsorber bed size. Empirical
design procedures based on adsorption equilibrium conditions are the easiest and
most common methods used to predict adsorber size and performance. These
methods will be discussed later.
Adsorption Forces — Physical and Chemical
The adsorption process is classified as either physical or chemical. The basic dif-
ference between physical and chemical adsorption is the manner in which the gas
molecule is bonded to the adsorbent. In physical adsorption the gas molecule is
bonded to the solid surface by weak forces of intermolecular cohesion. The
chemical nature of the adsorbed gas remains unchanged; therefore, physical
adsorption is a readily reversible process. In chemical adsorption a much stronger
bond is formed between the gas molecule and adsorbent. A sharing or exchange of
electrons takes place — as happens in a chemical bond. Chemical adsorption is not
easily reversible.
The forces active in physical adsorption are electrostatic in nature. These forces
are present in all states of matter: gas, liquid, and solid. They are the same forces
of attraction which cause gases to condense and real gases to deviate from ideal
behavior. This electrostatic force can be measured by the constant "a" in van der
Waals' equation describing nonideal gas behavior. Physical adsorption is sometimes
also referred to as van der Waals' adsorption. The electrostatic effect which pro-
duces the van der Waals' forces depends on the polarity of both the gas and solid
molecules. Molecules in any state are either polar or nonpolar depending on their
5-4
-------
chemical structure. Polar substances are those which exhibit a separation of
positive and negative charges within the compound. This separation of positive and
negative charges is referred to as a permanent dipole. Water is a prime example of
a polar substance. Nonpolar substances have both their positive and negative
charges in one center so they have no permanent dipole. Most organic compounds,
because of their symmetry, are nonpolar.
Physical, or van der Waals' adsorption can occur from three different effects: an
orientation effect, dispersion effects, or induction effects (Figure 5-3). For polar
molecules, attraction to each other occurs because of the orientation effect. The
orientation effect describes the attraction which occurs between the dipoles of two
polar molecules. The negative area of one is attracted to the positive area of the
other. An example of this type of adsorption would be the removal of water vapor
(polar) from an exhaust stream by using silica gel (polar).
Polar-Polar
wwv
Nonpolar — Nonpolar
Polar — Nonpolar
Figure 5-3. Physical forces causing adsorption.
The adsorption of a nonpolar gas molecule onto a nonpolar surface is accounted
for by the dispersion effect. The dispersion effect is based on the fact that although
nonpolar substances do not possess a permanent dipole, they do have a fluctuating
or oscillating dipole. Fluctuating dipoles are a result of momentary changes in elec-
tron distribution around the atomic nuclei. In a nonpolar substance, when two
fluctuating dipoles come close to one another, their total energy decreases, and
they fluctuate in phase with each other. Oscillating dipoles disperse light.
Consequently, this is where the name dispersion effect comes from. The adsorption
of organic vapors onto activated carbon is an example of nonpolar molecular
attraction.
5-5
-------
The attraction between a molecule with a permanent dipole (polar molecule)
and a nonpolar molecule is caused by the induction effect. A molecule with a per-
manent dlpole can induce or polarize a nonpolar molecule when they come in close
contact. The energy of this effect is determined by the polarizability of the non-
polar molecules. The induction effect is, however, very small when compared to
the orientation or dispersion effects. Therefore, adsorption systems use polar
adsorbents to remove polar contaminants and nonpolar adsorbents to remove non-
polar contaminants. This ensures that the intermodular forces of attraction
between the gas and solid will be greater than those between similar molecules in
the gas phase.
Chemical adsorption or chemisorption results from the chemical interaction
between the gas and the solid. The gas is held to the surface of the adsorbate by
the formation of a chemical bond. Adsorbents used in chemisorption can be either
pure substances or chemicals deposited on an inert carrier material. One example
is using pure iron oxide chips to adsorb H2S gases. Another example is using
activated carbon which has been impregnated with sulfur to remove mercury
vapors. 7
All adsorption processes are exothermic, whether adsorption occurs from
chemical or physical forces. In adsorption, molecules are transferred from the gas
to the surface of a solid. The fast-moving gas molecules lose their kinetic energy of
motion to the adsorbent in the form of heat.
In chemisorption, the heat of adsorption is comparable to the heat evolved from
a chemical reaction, usually over 10 kcal/mol. The heat given off by physical
adsorption is much lower, approximately 100 cal/mol, which is comparable to the
heat of condensation. Some additional general differences between physical
adsorption and chemisorption which make physical adsorption more desirable for
air pollution control are:
1. Molecules that are adsorbed by chemisorption are very difficult (and in some
cases, impossible) to remove from the adsorbent bed. Physically adsorbed
molecules can usually be removed by either increasing the operating
temperature or reducing the pressure of the adsorbent bed.
2. Chemisorption is a highly selective process. A gas molecule must be capable
of forming a chemical bond with the adsorbent surface for chemisorption to
occur. Physical adsorption occurs under suitable conditions in most gas-solid
systems. For industrial purposes specific solids are chosen which enhance the
rate of adsorption.
3. Chemisorption stops when all the reactive sites on the surface of the adsorbent
have reacted. Chemisorption forms only a monolayer of adsorbate molecules
on the surface. Because of van der Waals' forces, physical adsorption can
form multilayers of adsorbate molecules - one atop another.
4. The chemisorption rate increases with increasing temperature. For physical
adsorption the exact opposite is true: the rate decreases with increasing
temperature.
For these and other reasons, chemisorption is not used extensively in air pollution
control systems. A summary of the characteristics of physical versus chemical
adsorption is presented in Table 5-1.
5-6
-------
Table 5-1. Summary of characteristics of chemisorption and physical adsorption.
Chemisorption
Physical adsorption
Releases high heat
10,000 cal/mol
Forms a chemical compound
Desorption is difficult
Impossible adsorbate recovery
Releases low energy
100 cal/mol
Dipolar interaction
Desorption is easy
Easy adsorbate recovery
Adsorption Equilibrium Relationships
Most available data on adsorption systems is determined at equilibrium conditions.
Adsorption equilibrium is the set of conditions at which the number of molecules
arriving on the surface of the adsorbent equals the number of molecules that are
leaving. The adsorbent bed is said to be "saturated with vapors" and can remove
no more vapors from the exhaust stream. Equilibrium determines the maximum
amount of vapor that may be adsorbed at a given set of operating conditions.
Although a number of variables affect adsorption, the two most important in
determining equilibrium for a given system are temperature and pressure. Three
types of equilibrium graphs are used to describe adsorption systems: isotherm at
constant temperature; isobar at constant pressure; and isostere at constant amount
of vapors adsorbed.
Isotherm
The most common and useful adsorption equilibrium data is the adsorption
isotherm. The isotherm is a plot of the adsorbent capacity versus the partial
pressure of the adsorbate at a constant temperature. Adsorbent capacity is usually
given in weight percent expressed as gram of adsorbate per 100 grams of adsor-
bent. Figure 5-4 is a typical example of an adsorption isotherm for carbon
tetrachloride on activated carbon. Graphs of this type are used to estimate the size
of adsorption systems as illustrated in Example 5-1, following. Attempts have been
made to develop generalized equations which can predict adsorption equilibrium
from physical data. This is very difficult because adsorption isotherms take many
shapes depending on the forces involved. Isotherms may be concave upwards, con-
cave downwards or "S" shaped. To date, most of the theories agree with data only
for specific adsorbates-adsorbent systems and are valid over limited concentration
ranges.
5-7
-------
o
£ 100
o
o
U
u
f
c
u
rt
Q.
10
3.0
2.0
1.5
U 1.0
I I Ml
II 1 II
Mill
TTTT
' ' ' "Mill I I I
0.0001
0.001
' I III
0.1
I I I II
1.0
0.01
Partial pressure, psia
Source: Adapted from Technical Bulletin Calgon Corp.
Figure 5-4. Adsorption isotherm for carbon tetrachloride on activated carbon
Example 5-1*
A dry cleaning process exhausts a 15,000 scfm air stream containing 680 ppm car-
bon tetrachloride. Given Figure 5-4 and assuming the exhaust stream is at approx-
imately 140 °F and 14.7 psia, determine the saturation capacity of the carbon.
Solution:
In the gas phase, the mole fraction (Y) is equal to the percent by volume.
Y = % volume = 680 ppm
_ 680
106
= 0.00068
Obtaining the partial pressure:
p = YP
= (0.00068)(14.7psia)
= 0.01 psia
"This example problem is in English Units since Figure 5-4 needed to solve the problem is i
English Units.
in
5-8
-------
From Figure 5-4, at a partial pressure of 0.01 psia and a temperature of 140°F, the
carbon capacity is read as 30%.
This means that at saturation, 30 Ibs of vapor are removed per 100 Ibs of carbon
in the adsorber (30 kg/100 kg).
It must be noted that in practical applications adsorbers use more carbon than is
predicted at saturation to ensure that uncaptured vapors are not being exhausted
to the atmosphere. Example 5-2 will illustrate this point (page 5-18).
Polanyi Potential Theory
The most useful theory from an engineering design viewpoint, in trying to
predict adsorption isotherms, is the Polanyi potential theory (Cerny, 1970). The
Polanyi theory states that the adsorption potential is a function of the reversible
isothermal work done by the system. This is expressed in Equation 5-4 relating the
free energy of adsorption, AF, to the relative vapor pressure p°/p (Polanyi, 1914).
(Eq. 5-4) AF=-RTln(p°/p)
Where: AF= free energy of adsorption, cal/mol
T = absolute temperature, K
p° = vapor pressure of adsorbate, Pa
p = partial pressure of adsorbate, Pa
R = ideal gas constant
Polanyi observed that for any fixed amount of adsorbed material, AF is constant.
This relationship makes it possible to calculate the relative vapor pressures of a gas
in equilibrium with the adsorbent at a variety of temperatures given one measured
isotherm. If p°/p and T are known at a fixed amount of material adsorbed
(therefore AF) then from Equation 5-4, p°/p can be estimated for the same
amount adsorbed but at different temperatures. This relationship still requires that
experimental data be available and is only valid over a narrow temperature range:
from 60 to 130°F (EPA, 1972).
5-9
-------
Langmuir Theory
Another theory used to describe the isothermal adsorption process is the
Langmuir theory. Langmuir assumed that adsorption occurs in a unimolecular
layer so that a molecule can be adsorbed only if it strikes an uncovered area of the
adsorbent (Langmuir, 1918). The derivation of the Langmuir equation also
assumes that the forces acting in adsorption are the same as those for condensation
of vapors. The equation has no direct application to pollution control system
design, but is the basis for other useful equations. Brunauer, Emmett, and Teller
(1938) developed equations for multimolecular adsorption based on Langmuir's
ideas. A common form of the Brunauer, Emmett, and Teller (BET) equation is
given as:
(Eq. 5-5) E = _L_ + t^l\ (p/p°)
V(p*-p) VmC \VmCj(PP)
Where: V = volume of vapor adsorbed per unit weight of adsorbent
Vm = volume of vapor adsorbed per unit weight of adsorbent with a
layer one molecule thick
C = constant dependent on temperature
Equation 5-5 is used by manufacturers of adsorbents to report the surface area of
their products. A graph of [p/V(p°- p)] versus p/p° gives a straight line with a
slope of C- l/VmC and an intercept of l/VmC. From these relationships both C and
Vm can be computed. Just as with the Polanyi equation, the BET equation has
limitations to its use. Data from small-pore sized adsorbents deviate from the
straight line predicted by the BET equation. To date, there are no universally
accepted equations based on theory that are used to size adsorbers.
Isostere and Isobar
Two additional adsorption equilibrium relationships are the isostere and the isobar.
The isostere is a plot of the In p versus 1/T at a constant amount of vapor
adsorbed. Adsorption isostere lines are usually straight for most adsorbate-
adsorbent systems. Figure 5-5 is an adsorption isostere graph for the adsorption of
H2S gas onto molecular sieves. The isostere is important in that the slope of the
isostere corresponds to the heat of adsorption. This can also be seen from Equation
5-4. The last equilibrium relationship is the isobar. The isobar is a plot of the
amount of vapors adsorbed versus temperature at a constant pressure. Figure 5-6
shows an isobar line for the adsorption of benzene vapors on activated carbon.
Note that the amount adsorbed decreases with increasing temperature, which is
always the case for physical adsorption. Since these three relationships were
developed at equilibrium conditions, they depend on each other. By determining
one, such as the isotherm, the other two relationships can be determined for a
given system. In the design of a pollution control system, the adsorption isotherm is
by far the most commonly used equilibrium relationship.
5-10
------- An error occurred while trying to OCR this image.
-------
0 50 100 150 200 250 300 350 400
Temperature, °C
Figure 5-6. Adsorption isobar for benzene on carbon
(benzene at 10.0 mm Hg).
Adsorbent Materials
Several materials are used effectively as adsorbing agents. The most common
adsorbents used industrially are activated carbon, silica gel, activated alumina
(alumina oxide), and zeolites (molecular sieves). Adsorbents are characterized by
their chemical nature, extent of their surface area, pore size distribution, and par-
ticle size. In physical adsorption the most important characteristic in distinguishing
between adsorbents is their surface polarity. As discussed previously, the surface
polarity determines the type of vapors a particular adsorbent will have the greatest
affinity for. Of the above adsorbents, activated carbon is the primary nonpolar
adsorbent. It is possible to manufacture other adsorbing material having nonpolar
surfaces. Since their surface area is much less than that of activated carbon, they
are not used commercially. Polar adsorbents will preferentially adsorb any water
vapor that may be present in a gas stream. Since moisture is present in most pollu-
tant air streams, the use of polar adsorbents is severely limited for an air pollution
system. Therefore, in further discussion, the emphasis is placed on the use of
activated carbon, although some of the information is applicable to polar adsorp-
tion systems.
Activated Carbon
Activated carbon can be produced from a variety of feedstocks such as wood, coal,
coconut, nutshells, and petroleum-based products. The activation process takes
place in two steps. First, the feedstock is carbonized. Carbonization involves heating
5-12
-------
(usually in the absence of air) the material to a temperature high enough (600 °C)
to drive off all volatile material. Thus, carbon is essentially all that is left. To
increase the surface area, the carbon is then "activated" by using steam, air, or
carbon dioxide at higher temperatures. These gases attack the carbon and increase
the pore structure. The temperatures involved, the amount of oxygen present, and
the type of feedstock, all greatly affect the adsorption qualities of the carbon.
Manufacturers vary these parameters to produce activated carbons suitable for
specific purposes. In sales literature, the activity and retentivity of carbons are
based on their ability to adsorb a standard solvent, such as carbon tetrachloride
(CCU).
Because of its nonpolar surface, activated carbon is used to control emission of
organic solvents, odors, toxic gases, and gasoline vapors. Carbons used in gas phase
adsorption systems are manufactured in granular form, usually between 4x 6 to
4x 20 mesh in size. Bulk density of the packed bed can range from 0.08 to 0.5
g/cm3 (5 to 30 lb/fts) depending on the internal porosity of the carbon. Surface
area of the carbon can range from 600 to 1600 mVg (2.9x 10s to 7.8x 106 ftVlb).
This is equivalent to having the surface area of 2 to 5 football fields in one gram of
carbon.
Silica Gel
Silica gels are made from sodium silicate. Sodium silicate is mixed with sulfuric
acid, resulting in a jellylike precipitant from where the "gel" name comes. The
precipitant is then dried and roasted. Depending on the processes used in manufac-
turing the gel, different grades varying in activity can be produced. Silica gels have
surface areas of approximately 750 mVg (3.7 x 106 ftVlb). Silica gels are used
primarily to remove moisture from exhaust streams, but are ineffective at
temperatures above 260 °C (500 °F).
Molecular Sieves
Unlike the other adsorbents, which are amorphous (not crystalline) in nature,
molecular sieves have a crystalline structure. The pores are, therefore, uniform in
diameter. Molecular sieves can be used to capture or separate gases on the basis of
molecular size and shape. An example of this are refining processes which
sometimes use molecular sieves to separate straight chained paraffins from branched
and cyclic compounds. However, the main use of molecular sieves is in the removal
of moisture from exhaust streams. The surface area of molecular sieves range from
600 to 700 mVg (2.9x 10s to 3.4 X 106 ftVlb).
Aluminum Oxide (Activated Alumina)
Aluminum oxides are manufactured by thermally activated alumina or bauxite.
This is accomplished by heating the alumina in an inert atmosphere to produce a
porous aluminum oxide pellet. Aluminum oxides are not commonly used in air
pollution applications. They are primarily used for the drying of gases, especially
5-13
-------
under high pressures, and as support material in catalytic reactions. A prime
example is the impregnating of the alumina with platinum or palladium for use in
catalytic incineration. Activated alumina's surface areas can range from 200 to
300 mVg (0.98x 10s to 1.5x 106 ftVlb).
Pore Size Distribution
The physical properties of the adsorbent affect the adsorption capacity, rate, and
pressure drop across the adsorber bed. Table 5-2 summarizes these properties for
the above adsorbents. Since adsorption occurs at the gas-solid interface, the surface
area available to the vapor molecules determines the effectiveness of the adsorbent.
Generally, the larger the surface area, the higher the adsorbent's capacity.
However, the surface area must be available in certain pore sizes if it is to be effec-
tive as a vapor adsorber. Dubinin (1936) classified the pores in activated carbon as
micropores, transitional oores, or macropores. Micropores are openings whose radii
are 200 nanometers (20 A) or less. Pores larger than 2000 nanometers (200 A) are
macropores. Transitional pores are those with radii between 200 and 2000
nanometers.
Table 5-2. Physical properties of major types of adsorbents.
Adsorbent
Activated carbon
Activated alumina
Molecular sieves
Silica gel
Internal
porosity
(%)
55-75
30-40
40-55
70
Surface
area
(m'/g)
600-1600
200-300
600-700
750
Pore
volume
(cm'/g)
0.80-1.20
0.29-0.37
0.27-0.38
0.40
Bulk dry
density
(g/cm1)
0.35-0.50
0.90-1.00
0.80
0.70
Mean pore
diameter
(mn)
150-200
180-200*
30-90
220
"The 150-200 nanometer average is for the micropores only; since 95% of the surface area is
associated with them.
Most gaseous air pollutant molecules are in the 40 to 90 nanometer size range. If
a large portion of an adsorbent's surface area is in pores smaller than
40 nanometers, many contaminant molecules will be unable to reach these active
sites. Figure 5-7 illustrates molecule movement in the pores. In addition, the larger
pore sizes (macro and transitional) contribute little to molecule capture. The vapor
pressure of the contaminant in these larger areas is too low to be effectively
removed. These larger pores serve mainly as passageways to the smaller micropore
area where the adsorption forces are strongest. Adsorption forces are strongest in
pores that are not more than approximately twice the size of the contaminant
molecule. These strong adsorption forces result from the overlapping attraction of
the closely spaced walls.
5-14
-------
Macropore
Molecule
blocking pore
Area unavailable
for adsorption
Figure 5-7. Molecular screening in pores of activated carbon.
Another phenomenon, capillary condensation usually only occurs in the
micropores. Capillary condensation occurs when multilayers of adsorbed contami-
nant molecules build up from both sides of the pore wall, totally packing the pore
and condensing in the pore. The amount of contaminant that is removed increases
since additional molecules condense on the surface of the liquid which has formed.
Contaminant molecules can also be removed at lower vapor pressures (more dilute
concentrations) since capture forces are now acting from three sides instead of just
one. However, desorption is not as complete if capillary condensation occurs, since
the forces that hold a liquid together are much stronger than the physical adsorp-
tion forces.
Air pollution control involves removing contaminant vapors at low partial
pressures. Therefore, the micropore structure of an adsorbent plays an important
role in determining the overall efficiency. Another reason for the wide use of
activated carbon is that 90 to 95% of its surface area is in the micropore size
range.
5-15
-------
Dynamic Adsorption Process
The movement of vapors through an adsorbent bed is often referred to as a
dynamic process. The term dynamic refers to motion both in the movement of air
through the adsorbent bed and change in vapor concentration as it moves through
the bed. There are a variety of configurations in which the contaminant air stream
and adsorbent are brought into contact. The most common configuration is to pass
the air stream down through a fixed volume or bed of adsorbent. Figure 5-8
illustrates how adsorption (mass transfer) occurs as vapors pass down through the
bed.
Mass transfer
zone
Saturated
bed
r ^^^•••••^•••^••^^^ ^*t^^^^^^^xmi^
TT
O 3
U .5
•hJ
• Breakpoint
in Volume of effluent treated (time)
Figure 5-8. Breakthrough curve.
The gas stream containing the pollutant, at an initial concentration, c,, is passed
down through a deep bed of adsorbent material which is free of any contaminant.
Most of the contaminant is readily adsorbed by the top portion of the bed. The
small amount of contaminant that is left is easily adsorbed in the remaining section
of the bed. The effluent from the bottom of the bed is essentially pollutant free,
denoted at d.
After a period of time the top layer of the adsorbent bed becomes saturated with
contaminant. The majority of adsorption (approximately 95%) now occurs in a
narrow portion of the bed directly below this saturated section. The narrow zone of
adsorption is referred to as the mass transfer zone (MTZ). As additional contami-
nant vapors pass through the bed, the saturated section of the bed becomes larger
and the MTZ moves further down the length of the adsorber. The actual length of
5-16
-------
the MTZ remains fairly constant as it travels through the adsorbent bed. Addi-
tional adsorption occurs as the vapors pass through the "unused" portion of the
bed. The effluent concentration at cz is essentially still zero since there is still an
unsaturated section of the bed.
Finally, when the lower portion of the MTZ reaches the bottom of the bed, the
concentration of contaminant in the effluent suddenly begins to rise. This is
referred to as the breakthrough point —where untreated vapors are being exhausted
from the adsorber. If the contaminated air stream is not switched to a fresh bed,
the concentration of contaminant in the outlet will quickly rise until it equals the
initial concentration, illustrated at point c4.
In most air pollution control systems even trace amounts of contaminants in the
effluent stream are undesirable. To achieve continuous operation, adsorbers must
be either replaced or cycled from adsorption to desorption before breakthrough
occurs. In desorption or regeneration, the contaminant vapors are removed from
the used bed in preparation for the next cycle. Most commercial adsorption systems
are the regenerable type.
In regard to regenerable adsorption systems, three important terms are used to
describe the capacity of the adsorbent bed. All the capacities are measured in Ib of
vapor per Ib of adsorbent. First, the breakthrough capacity is defined as the
capacity of the bed at which unreacted vapors begin to be exhausted. The satura-
tion capacity is the maximum amount of vapors that can be adsorbed per unit
weight of carbon. (This is the capacity read from the adsorption isotherm). The
working capacity is the actual amount of adsorbent used in an adsorber. The
working capacity is a certain fraction of the saturation capacity. Working
capacities can range from 0.1 to 0.5 of the saturation capacity. (Note: a smaller
capacity increases the amount of carbon required.) This fraction is set by the
designer for individual systems by balancing the cost of carbon and adsorber opera-
tion versus preventing breakthrough allowing for an adequate cycle time.
Another factor in determining the working capacity is that it is uneconomical to
desorb all the vapors from the adsorber bed. The small amount of residual vapors
left on the bed is referred to as the heel. This heel accounts for a large portion of
the difference between the saturation and the working capacity. In some cases the
working capacity can be estimated by assuming it is equal to the saturation capa-
city minus the heel (Turk, 1977). The following example illustrates one method of
estimating the working capacity. In all the examples in this manual and the
accompanying workbook, a design factor of 0.5 of the saturation capacity is used.
This is the same as assuming the working amount of carbon is twice the amount
required at saturation.
5-17
-------
Example 5-2
Assume the same conditions as stated in Example 5-1. Estimate the amount of car-
bon that would be required if the adsorber were to operate on a 4-hour cycle. The
molecular weight of CC14 is 154 Ib/lb mole.
Solution:
From Example 5-1 we know that the carbon used will remove 30 Ibs of vapor for
every 100 Ibs of carbon at saturation conditions.
First compute the flow rate of CC14:
Qfc;4 =15,000 scfm x 0.00068
= 10.2 scfm CC14
Converting to pounds per hour:
10.2 £ x lb mole x 1541b x 60min =262 5 lb CC1*
mm 359 ft3 lb mole hr ' hr
For a 4-hour cycle there are: 4 X 262.5= 1050 lb CC14.
The amount of carbon (at saturation) required:
1050 lb CCU X u, = 3500 lb carbon
30 lb CC14
The actual amount of carbon required can be estimated by doubling the amount
needed at saturation.
2 X 3500 = 7000 lb (3182 kg) carbon per 4-hour cycle per adsorber.
Note: this gives only a rough estimate of the amount of carbon needed.
Factors Affecting Adsorption
A number of factors or system variables influence the performance of an adsorp-
tion system. These variables and their effects on the adsorption process are dis-
cussed in the following section.
Temperature
For physical adsorption processes, the capacity of an adsorbent decreases as the
temperature of the system increases. Figure 5-9 illustrates this concept. As the
temperature increases the vapor pressure of the adsorbate increases, raising the
energy level of the adsorbed molecules. Adsorbed molecules now have sufficient
energy to overcome the van der Waals' attraction and migrate back to the gas
phase. Molecules already in the gas phase tend to stay there due to their high
5-18
-------
vapor pressure. As a general rule, adsorber temperatures are kept below 55°C
(130°F) to ensure adequate bed capacities. Temperatures above this limit can be
avoided by cooling the exhaust stream that is to be treated.
u
a
O,
u
c
(9
U
Temperature
Figure 5-9. Carbon capacity vs. temperature.
Adsorption is an exothermic process with the heat released for physical adsorp-
tion approximately equal to the heat of condensation. At low concentrations (below
100 ppm) the heat release is minimal and is quickly dissipated by the air flow
through the bed. At higher concentrations (approximately 5000 ppm) considerable
heating of the bed can occur, which if not removed can cause the adsorber effi-
ciency to rapidly decrease. In addition, granular carbon is a good insulator, which
inhibits heat dissipation from the interior of the bed. In some cases, especially
ketone recovery, the temperature rise can cause auto-ignition of the carbon bed.
Monitoring of bed temperatures and leaving the bed slightly wet after steam
regeneration are techniques used to avoid bed fires.
Pressure
Adsorption capacity increases with an increase in the partial pressure of the vapor.
The partial pressure of a vapor is proportional to the total pressure of the system.
Any increase in pressure will increase the adsorption capacity of a system (see
Figure 5-4). The increase in capacity occurs because of a decrease in the mean free
path of vapors at higher pressures. Simply, the molecules are packed more tightly
together. More molecules have a chance to "hit" the available adsorption sites,
increasing the number of molecules adsorbed.
5-19
-------
Gas Velocity
The contact or residence time between the contaminant stream and adsorbent is
determined by the gas velocity through the adsorber. The residence time directly
affects capture efficiency. The slower the contaminant stream flows through the
adsorbent bed, the greater the probability of a contaminant molecule hitting an
available site. Once a molecule has been captured it will stay on the surface until
the physical conditions of the system are changed. To achieve 90% + capture
efficiency most carbon adsorption systems are designed for a maximum air flow
velocity of 30 m/min (100 ft/min) through the adsorber. A lower limit of at least
6 m/min (20 ft/min) is maintained to avoid flow distribution problems, such as
channeling.
Gas velocity through the adsorber is a function of the diameter of the adsorber
for a given volume of contaminant gas. By specifying a maximum velocity through
the adsorber, the minimum diameter is also specified. For example, if 300 rnVmin
of contaminant gas is to be treated, and the maximum velocity through the
adsorber is to be 30 m/min, then the adsorber must have a cross sectional area of
at least 10 m2.
The gas flow rate through the adsorber also affects the pressure drop. Increasing
the flow rate increases the pressure drop. Within the above stated maximum and
minimum flow rates, the allowable pressure drop usually dictates the required
tower diameter and flow rate. The pressure drop across the bed also depends on
the depth of adsorbent. This will be discussed in the following section.
Bed Depth
Providing a sufficient depth of adsorbent is very important in achieving efficient
gas removal. If the adsorber bed depth is shorter than the required mass transfer
zone, breakthrough will immediately occur rendering the system ineffective. Com-
puting the length of the MTZ is very difficult since it depends upon six factors: the
adsorbent particle size, gas velocity, adsorbate concentration, fluid properties of the
gas stream, temperature, and pressure of the system. The MTZ can be estimated
from experimental data using Equation 5-6 (Kovach, 1978). To obtain the
necessary data, vendors will usually test a small portion of the exhaust stream on a
pilot adsorber column.
(Eq. 5-6) MTZ = D (l - ^
1 - Xs \ Cs
Where: D = bed depth, m
Cfl = breakthrough capacity, %
Gy = saturation capacity, %
Xs = degree of saturation in the MTZ, % (usually assumed to be 50%)
MTZ = length of MTZ, m
The above equation is used mainly as a check to ensure that the proposed bed
depth is longer than the MTZ. Actual bed depths are usually many times longer
than the length of the MTZ. The additional bed depth allows for adequate cycle
times. Equation 5-6 can be rearranged to solve for the breakthrough capacity:
5-20
-------
(Eq. 5-7)
_ _ (Xj)(Q)(MTZ) + CXD - MTZ)
CB-----
The total amount of adsorbent required is usually determined from the adsorp-
tion isotherm, as illustrated in Example 5-2. Once this has been set, the bed depth
can then be estimated by knowing the tower diameter and density of the adsorbent.
Example 5-3 illustrates how this is done. Generally, the adsorbent bed is sized to
the maximum length allowed by the pressure drop across the bed. Data on the
pressure drop per meter of bed depth for typical carbons is presented in Figure
5-10 (Turk, 1977). The pressure drop per meter of bed depth is plotted versus the
gas flow rate, with the carbon mesh size as a parameter. From the figure, an
adsorber with a flow rate of 40 cm/s (80 ft/min) using 4 X 10 mesh carbon will
have a pressure drop of approximately 5 kPa per meter (6 in. H2O per foot) of bed
depth. Therefore, if the pressure drop across the bed is limited to 4.5 kPa (18 in.
H2O) then the total bed depth should not exceed 0.9 meters (3 ft).
Linear velocity, cm/s
10 20
40 50
20 30 40 50
Linear velocity, ft/min
100
Source: From Air Pollution, 3rd ed., Vol IV, Engineering Control of Air Pollution,
Chapter 8-Adsorption by Amos Turk, A.C. Stern editor. ©1977. Used with the
permission of Academic Press, Inc.
Figure 5-10. Pressure drop vs. flow rate through granular carbon beds.
5-21
-------
Humidity
As stated previously activated carbon will preferentially adsorb nonpolar hydro-
carbons over polar water vapor. The water vapor molecules in the exhaust stream
exhibit strong attractions for each other rather than the adsorbent. At high relative
humidities, over 50%, the number of water molecules increases such that they
begin to compete with the hydrocarbon molecules for active adsorption sites This
reduces the capacity and the efficiency of the adsorption system. Exhaust streams
with humidities greater than 50% may require installation of additional equipment
to remove some of the moisture. Coolers to remove the water are one solution
Dilution air with significantly less moisture in it than the process stream has also
been used. Also, the contaminant stream may be heated to reduce the humidity as
long as the increase in temperature does not greatly affect adsorption efficiency.
Contaminants
In addition to humidity; paniculate matter, entrained liquid droplets, and organic
compounds which have high boiling points can also reduce adsorber efficiency if
present in the air stream. Any micron-sized particle of dust or lint which is not
filtered can cover the surface of the adsorbent. This greatly reduces the surface
area of the adsorbent available to the gas molecule for adsorption. Covering of
active adsorption sites by an inert material is referred to as blinding or
deactzvation. To avoid this situation almost all industrial adsorption systems are
equipped with some type of paniculate matter removal device.
Entrained liquid droplets can also cause operational problems. Liquid droplets
that are nonadsorbing act the same as paniculate matter. The liquid covers the
surface, blinding the bed. If the liquid is the same as the adsorbate, high heats of
adsorption occur. This is especially a problem in activated carbon systems where
liquid organic droplets carried over from the process can cause bed fires from the
heat released. Some type of entrainment separator may be required when liquid
droplets are present.
For activated carbon systems, other contaminants are high boiling point organic
compounds, usually in excess of 260°C (500°F). High boiling point (high molecular
weight) compounds have such an affinity for the carbon that it is extremely dif-
ficult to remove them by standard desorption practices. These compounds also tend
to react chemically on the carbon surface forming solids or polymerization products
which are extremely difficult to desorb. Loss of carbon activity in this manner is
called chemical deactivation.
Adsorbent Regeneration Methods
Periodic replacement or regeneration of the adsorbent bed is mandatory in order to
maintain continuous operation. When the adsorbate concentration is high and/or
cycle time is short (less than 12 hours) replacement of the adsorbent is not feasible
5-22
-------
and in-situ regeneration is required. Regeneration is accomplished by reversing the
adsorption process, usually by increasing the temperature or decreasing the
pressure. Commercially, four methods are used in regeneration:
Thermal swing: The bed is heated so that the adsorption capacity is reduced to a
lower level. The adsorbate leaves the surface of the carbon and is removed from
the vessel by a stream of purge gas. Cooling must be provided before the sub-
sequent adsorption cycle begins. Steam regeneration is an example of thermal
swing regeneration.
Pressure swing: The pressure is lowered at a constant temperature to reduce the
adsorbent capacity.
Inert purge gas stripping: The stripping action is caused by an inert gas that
reduces the partial pressure of the contaminant in the gas phase, reversing the con-
centration gradient. Molecules migrate from the surface into the gas stream.
Displacement cycle: The adsorbates are displaced by some preferentially
adsorbed material. This method is usually a last resort for situations in which the
adsorbate is both valuable and is heat sensitive, and for which pressure swing
regeneration is ineffective (Bethea, 1978).
Table 5-3 compares the effectiveness of the various regeneration methods (Wood,
1964). As can be seen from this table, steam regeneration was most effective for the
test conditions. This is also true for most industrial applications.
Table 5-3. Regeneration of one pound of activated carbon loaded with 20% ether.
Regeneration method
Thermal swing
Pressure swing
Combination
Thermal swing
Regeneration conditions
Heating at 100°C (212°F) for 20 min
Vacuum of 50 mm Hg at 20 °C (68 °F) for 20 min
Gas circulation at 130°C (266 °F) for 20 min
Direct steam at 100°C (212°F) for 20 min
Expelled ether
(%)
15
25
45
98
are:
Thermal Swing—Steam Stripping
Because it is simple and relatively inexpensive, steam stripping is the most common
desorption technique. Several additional advantages to using steam for desorption
At high pressure, the steam's temperature (100°C) is high enough to desorb
most solvents of interest without damaging the carbon or the desorbed vapors.
Desorbed vapors can be polymerized or cracked, sometimes forming
undesirable compounds.
Steam readily condenses in the adsorber bed releasing its (the steam) heat of
condensation, aiding in desorption.
Many organic compounds can be easily separated and recovered from the
effluent stream by condensation or distillation.
5-23
-------
• Residual moisture in the bed is removed easily by a stream of cool dry air
(either pure or process effluent).
• Steam is a more concentrated source of heat than hot air so it is very effective
in raising the temperature of the adsorber bed very quickly.
The amount of steam required for regeneration depends on the adsorbate
loading of the bed. The longer a carbon bed is steamed, the more adsorbate will
be desorbed. It is usually not cost effective to try to desorb all of the adsorbed
vapors from the bed. Acceptable working capacities can be achieved by using less
steam and leaving a small portion of adsorbate in the bed. During the initial
heating period no vapors are desorbed. This is because a fixed amount of steam is
first required to raise the temperature of the cold bed to the desorption
temperature. After this initial period a substantial amount of adsorbate vapor is
released, until a plateau is reached. The plateau represents the optimum steam
requirement, usually in the range 0.25 to 0.35 kg of steam/kg of carbon (Parmele,
1979). In these systems, steam is usually supplied at pressures ranging from 21 to
103 kPa (3 to 15 psig).
A typical two bed adsorption system is shown in Figure 5-11. Regeneration steam
usually passes up through the bed countercurrent to the flow of solvent laden
vapors. Since the bed is switched to the desorption mode before breakthrough, the
outlet end of the bed remains adsorbate free, providing a safety margin for sub-
sequent cycles. The steam usage can range anywhere from 0.3 to 10 kg of steam
per kg of solvent removed.
Some disadvantages are associated with steam regeneration. Problems arising
are:
• The effluent from the condenser could pose a water pollution problem unless
the condensate is sent to a waste water treatment facility.
• Some organic compounds are subject to hydrolysis and/or other reactions with
water which may produce corrosive substances. Corrosive substances can
greatly reduce the life of the adsorption equipment unless expensive corrosive
resistant materials are used.
• A hot wet carbon bed will not effectively remove organic vapors. Cooling and
drying of the bed may be needed to ensure adequate removal efficiencies at
the beginning of a subsequent cycle.
5-24
------- An error occurred while trying to OCR this image.
-------
Pressure Swing— Vacuum Desorption
Pressure swing or vacuum desorption has one primary advantage over thermal
desorption. Desorption is accomplished by a change in pressure rather than
temperature so no time is required to initially heat up or cool the carbon bed. This
adiabatic (no change in temperature) pressure swing allows the bed to be in the
adsorbing cycle longer. Units may also be sized smaller since there is no increase in
air volume due to heating of the bed. Both of these conditions allow for higher
throughputs or smaller adsorber designs than can be accommodated by thermal
swing desorption systems. In addition, the desorbed vapors may be recovered
directly without the need for additional downstream processing equipment.
The principle disadvantages of a pure pressure swing cycle are the high
operating and construction costs. In pressure/vacuum systems the adsorber vessel
and valving must be constructed of materials capable of withstanding vacuums of
9.5 kPa (28 in. Hg). Unless the adsorber is initially operated at elevated pressures
(so that the pressure swing can be accomplished by reducing the vessel to
atmospheric pressures) a vacuum producing system is required. Vacuum systems
that operate cyclically may require more operating attention than other regenera-
tion systems. To be effective, pressure regeneration systems must be designed so
that a small decrease in pressure will result in a drastic shift in the direction of
mass transfer.
Adsorption Control Systems
Adsorption control systems can be classified as either regenerable or non-
regenerable. Nonregenerable systems are normally used to control exhaust streams
with low pollutant concentrations, below 1.0 ppm. Generally these pollutants are
highly odorous or to some degree toxic. When these systems reach the
breakthrough point the bed is taken off stream and replaced with a fresh bed. The
used carbon can then be sent back to the manufacturer for reactivation.
Regenerable systems are used for higher pollutant concentrations such as in solvent
recovery operations. Once the bed reaches the breakthrough point in a regenerable
system, the pollutant vapors are directed to a second bed while the first has the
vapors desorbed.
Nonregenerable Adsorption Systems
Nonregenerable adsorption systems are manufactured in a variety of configura-
tions. Bed areas are sized to control the air flow through them at between 6.0 to 18
m/min (20 to 60 ft/min). They usually consist of thin adsorbent bed depth,
ranging in thickness from 1.25 to 10.0 cm (0.5 to 4.0 inches). These thin beds have
a low pressure drop, normally below 62 Pa (0.25 in. H2O) dependent on the bed
thickness, gas velocity, and particle size of the adsorbent. Service time for these
units can range from 6 months for "heavy" odor concentrations to up to 2 years for
5-26
-------
trace concentrations or intermittent operations (EPA, 1973). They are used mainly
as air purification devices for small air flows in offices, laboratory exhaust, and
other small exhaust streams.
The shapes of these thin bed adsorbers are flat, cylindrical, or pleated. The
granules of activated carbon are retained by porous support material, usually per-
forated sheet metal. An adsorber system usually consists of a number of retainers or
panels placed in one frame. Figure 5-12 shows a nine panel thin bed adsorber. The
panels are similar to home air filters except that instead of containing steel wool
they contain activated carbon as the filter. Figure 5-13 illustrates a pleated cell
adsorber. The pleated cell is one continuous retainer of activated carbon, rather
than individual panels. The cylindrical canisters (Figure 5-14) are usually small
units designed to handle low flow rates of approximately 0.12 m/s (25 cfm). Cylin-
drical canisters are made of the same materials as the panel and pleated adsorbers
except their shape is round rather than square. Panel and pleated beds are dimen-
sionally about the same size, normally 0.6 meters square (2 ft by 2 ft) with the car-
bon depth ranging from 0.2 to 0.6 m (8 in. to 2 ft). Flat panel beds are sized to
handle higher exhaust flow rates, approximately 9.4 m/s (2000 cfm), while pleated
beds are limited to flow rates of 4.7 m/s (1000 cfm). Typical flow rate values are
listed in Table 5-4.
Carbon
panel
Figure 5-12. Thin bed adsorber: nine cell system.
5-27
-------
Figure 5-13. Pleated thin bed.
•Activated carbon
Figure 5-14. Canister.
Table 5-4. Adsorption filters.
Filter shape
Multiple panel cell
Pleated cell
Cylindrical canister
Size
-0.6 m2
-0.6 m*
-0.002 m diameter
-0.005 m length
Flow rate
9.4 m/s
4.7 m/s
0.12 m/s
5-28
-------
In addition to thin bed systems, thick bed nonregenerable systems are also
available. One system that can be used is essentially just a 55-gallon drum. The
bottom of the drum is filled with gravel to support a bed of activated carbon
weighing approximately 330 kg (150 Ib). A typical unit is shown in Figure 5-15.
These units are used to treat small flow rates (0.5 m/s or 100 cfm) from laboratory
hoods, chemical storage tank vents, and chemical reactors.
Activated carbon
Support material
Figure 5-15. Canister.
Regenerable Adsorption Systems
A large regenerable adsorption system can be categorized as a fixed, moving or
fluidized bed. The name refers to the manner in which the vapor stream and
adsorbent are brought into contact. The choice of a particular system depends on
the pollutants to be controlled and the recovery requirements. The most common
adsorption system for controlling air pollutants is the fixed carbon bed. These
systems are used to control a variety of organic vapors and are usually regenerated
by direct steaming of the bed. The organic compounds may be recovered by con-
densing the exhaust from the regeneration step and separating out the water and
solvent.
5-29
-------
Fixed Bed Adsorbers
Fixed bed adsorption systems generally involve multiple beds. One or more beds
treat the process exhaust while the other beds are either being regenerated or
cooled. A typical three bed adsorption system is shown in Figure 5-16. The solvent-
laden air stream is first pretreated to remove any solid or liquid particles which
could blind the carbon bed and decrease its efficiency. The solvent-laden air
stream then usually passes down through the fixed carbon bed. Upward flow
through the bed is usually avoided (unless flow rates are low (< 500 cfm) to
eliminate the risk of entraining carbon particles in the exhaust stream).
Pretreatment
Outlet
Condenser
Regenerating
steam
Figure 5-16. Three bed system.
After a predetermined length of time, referred to as the cycle time, the solvent-
laden air stream is directed to the second adsorber by a series of valves. Steam is
then injected into the first bed to remove the adsorbed vapors. The steam and
desorbed vapors are then usually sent to a recovery system. If the solvents are
immiscible in water, they can be separated by condensing the exhaust and
decanting off the solvent. If the solvents are miscible in water, distillation may be
required. Before the first adsorber is returned to service, cooling and drying of the
carbon should be provided. This will ensure against immediate breakthrough
occurring from the "hot, wet" carbon bed. This can be accomplished by venting
the solvent-laden air stream through the hot, wet adsorber, then to the on-line
adsorber to maintain a high removal efficiency.
Regenerable fixed carbon beds are usually between 0.3 and 1.2 m (1 to 4 ft)
thick. The maximum adsorbent depth of 1.2 m is based on pressure drop
considerations (Vic Mfg. Co.). Superficial gas velocities through the adsorber range
from 6.0 to 30.0 m/min (20 to 100 ft/min) with 30.0 m/min being a maximum
5-30
-------
permissible flow rate. Pressure drops normally range from 750 to 3730 Pa (3 to 15
in. H4O) depending on the gas velocity, bed depth and carbon particle size
(Bethea, 1978). For specific applications, graphs similar to that in Figure 5-11 are
supplied by the carbon manufacturer to compute the pressure drop.
The two types of fixed bed adsorbers are distinguished by bed orientation in
relation to air flow. The first is referred to as a vertical flow adsorber. The bed
length is vertical as is the direction of air flow. The air stream usually flows
downward. This system is shown in Figure 5-16. These units are suited to handle
flows up to 1.4 to 2.0 mVs (3000 to 4000 cfm) per adsorber. Figure 5-17 shows a
three bed vertical system used to recover 34 kg/m (75 Ibs/hr) of trichloroethylene
from a vapor degreasing operation. Each vessel is 1.2 m (48 in.) in diameter,
contains 255 kg (560 Ibs) of carbon approximately 10 cm (4 in.) deep and handles
1.7 mVs (3500 scfm) air flow. One vessel is being desorbed each hour using 102 kg
(225 Ibs) of steam at 103 kPa (15 psig). The pressure drop across the system is
approximately 3 kPa (12 in. H2O) (Vic Mfg. Co. with permission).
Figure 5-17. Three bed vertical system.
5-31
-------
For larger flow rates, horizontal flow adsorbers are used. Structurally they are
more suited to handle the larger air volumes. In horizontal flow units, the bed
length is horizontal as is the direction of the incoming air stream. The air stream
flows across the bed and down. This system is illustrated in Figure 5-18. Adsorbers
of this type are manufactured as a package system capable of handling flow rates
up to 1150 mVs (40,000 cfm). Larger units must be engineered and fabricated for
the specific application. Figure 5-19 shows a three bed horizontal system used to
recover 1180 kg/hr (2600 Ib/hr) of toluene from a rotogravure operation. Each
vessel is 7 m (22 ft) long and 3 m (10 ft) in diameter and contains 9000 kg (20,000
Ibs) of carbon packed to a 1 m (3 ft) bed depth. The system handles 21 mVs
(44,000 scfm) air flow with a pressure drop of 6 kPa (24 in. H2O). One vessel is
desorbed each 45 minutes using 3600 kg/hr (8000 Ibs/hr) of steam at 103 kPa
(15 psig) pressure (Vic Mfg. Co. with permission).
Steam and
vapor outlet
Carbon bed
Steam inlet
Figure 5-18. Horizontal bed.
5-32
-------
Figure 5-19. Three bed horizontal system.
Moving Bed Adsorbers
Moving bed systems can also be used to obtain a higher degree of utilization of the
carbon bed than is possible with a fixed bed. In moving bed systems, the solvent-
laden vapor stream passes only through the unsaturated portion of the carbon bed.
This reduces the distance (thus pressure drop) the air stream travels through the
bed.
One design of a moving (rotary) bed system is illustrated in Figure 5-20 (Sutcliffe
Speakman Co., 1963). The device consists of four cylinders which are in constant
rotation. The granular carbon is held in place between two cylinders made of steel
screening or perforated sheet metal. This bed is placed between inner and outer
cylinders which are impervious to air flow except at slots near their ends.
The slots on the outer cylinder act as solvent-laden inlets. They permit the air
stream to pass into the annular section where the carbon is located. The solvent-
laden air stream passes through the carbon bed and purified air exits out the inner
slots. The carbon bed is broken into sections. The cylinders rotate such that when
the proper degree of saturation is reached the bed is desorbed. Desorption occurs
by injecting steam in through the slots on the inner cylinder. Steam and desorbed
vapors exit through the slots on the outer cylinder. During each rotation of the
cylinder, each segment of the carbon bed undergoes both an adsorption and
desorption cycle.
Because of the continuous adsorbing and desorbing process, bed utilization is
improved. The air stream is no longer required to pass through the top, saturated
portion; or the bottom, idle portion of the bed. The air stream passes only through
the active, mass transfer portion of the bed. Therefore, shorter and more compact
beds may be used which reduce the pressure drop. The disadvantages are wear on
moving parts and maintaining air tight seals on moving parts.
5-33
-------
Vapor inlet
Outer shell
Outer screen
Inner screen
Inner shell
Regenerative steam
inlet
Activated carbon
et 1 i
Steam and
vapor outlet
Vapor
inlets
Steam and
vapor outlet
'\
Interior
outlets
Figure 5-20. Rotary bed system.
Fluidized Bed Adsorbers
A fluidized bed adsorption system operates in the same physical manner as a tray
scrubber. Instead of liquid flowing down the column from tray to tray, granular
activated carbon is used. Figure 5-21 shows one recently developed fluidized bed
system which is being marketed by Union Carbide. The solvent-laden air stream is
introduced at the middle of the tower. Then it passes up through the tower
fluidizing the activated carbon in a series of trays. The carbon then flows down
through the vessel from tray to tray until it reaches the desorption section.
Regeneration is accomplished in the bottom half of the vessel and the activated
carbon is air conveyed back to the top of the tower.
5-34
-------
Clean air out
Perforated trays
Air lift
blower
Desorber
heating
section
4*^ Nitrogen recycle
blower
Recovered
solvent
Figure 5-21. Fluidized bed adsorber.
As with the moving bed, the fluidized bed also provides continuous operation
and more efficient utilization of the adsorbent. The need for multiple vessels is
eliminated, which can greatly reduce the cost of the system. Gas velocities around
1 m/s (196 ft/min) are needed to fluidize the bed. These are 2 to 4 times the
velocities achieved in fixed bed systems. This allows for use of a much smaller
vessel for comparable air flow and helps to achieve uniform gas distribution.
5-35
-------
The main disadvantage with fluidized bed adsorption systems is the high attrition
(wear) losses of the granular activated carbon. Recently a new "beaded" activated
carbon was developed in Japan. The beaded shape is inherently stronger and has
better fluidity properties than granular carbon. The beaded carbon has been used
in a number of installations (mostly in Japan) and is reported to reduce the attri-
tion losses to 2 to 5% per year as compared to 10% for fluidized granular carbon
(Union Carbide).
The following example illustrates the use of the principles and general rules of
practice discussed in this chapter in designing an adsorption system.
Example 5-3
A solvent degreaser is designed to recover toluene from a 3.78 mVs (8000 acfm) air
stream at 25 °C (77 °F) and atmospheric pressure. The company is planning to use
a two bed carbon adsorption system with a cycle time of 4 hours. The maximum
concentration of toluene is kept below 50% of the lower explosive limit for safety
purposes. Given Figure 4-4 in the 415 Student Workbook, the adsorption isotherm
for toluene, and the additional operational data, estimate:
1. the amount of carbon required for a 4-hour cycle
2. square feet of surface area required based on a 0.508 m/s (100 fpm) maxi-
mum velocity
3. depth of the carbon bed
Given: LEL for toluene = 1.2%
molecular weight of toluene = 92.1 kg/kg mol
carbon density = 480 kg/m3 (30 lb/ft3)
Solution:
1. To compute amount of carbon required, first, calculate the toluene flow rate.
(3.78 mVs)(50%)(1.2%) = 0.023 mVs toluene
To determine the saturation capacity of the carbon, calculate the partial
pressure of toluene at the adsorption conditions.
p = YP
.
1(14.7 psia)
F '
3.78 mVs
= 0.089 psia
From Figure 4-4 in the 415 Student Workbook, the saturation capacity of the
carbon is 40% or 40 kg toluene per 100 kg of carbon at 0.089 psia.
The flow rate of toluene is:
0.023
s / \22.4 m3/\350 K/\kg mol
= 0.074 kg/ s of toluene
5-36
-------
The amount of carbon at saturation for a 4-hour cycle is:
0.074 IS toluene100 aron
v s A 40 kg toluene A hr r >
= 2664 kg of carbon
The working charge of carbon can be estimated by doubling the saturation
capacity. Therefore, the working charge is:
working charge = (2)(2664 kg of carbon)
= 5328 kg of carbon
2. The square feet of superficial surface area is the surface area set by the maxi-
mum velocity of 0.508 m/s (100 fpm) through the adsorber.
The required surface area is:
maximum velocity
3.78 m'/s
0.508 m/s
= 7.44 m2
For a horizontal flow adsorber this would correspond to a vessel approximately
2 m (6.6 ft) in diameter and 4 m (13.1 ft) in length to give 8.0 m2 (87 ft2)
surface area. This would supply more than the required area.
The flow rate is too high to be handled by a single vertical flow adsorber. An
alternative would be to use three vessels, two adsorbing while one is being
regenerated. Each vessel then must be sized to handle 1.89 mVs (4000 acfrn).
The area required for a limiting velocity of 0.508 m/s is:
Area to handle half flow = L89 mVs
0.508 m/s
= 3.72m2
This cross sectional area corresponds to a vessel diameter of:
= 2.18m (7 ft)
5-37
-------
3. The volume that the carbon would occupy in the horizontal bed system is:
Volume of carbon = weight X
density
= 5328 kg X m
480kg
= 11.1 m3
Note: for the three bed vertical system the volume of each bed would be half
this or 5.55 m3.
4. The depth of carbon for the horizontal bed is given by:
Depth of carbon = Volume of carbon
Cross sectional area of adsorber
11.1 m3
7.44 m2
= 1.49 m
Note: the depth for the three bed vertical system is the same, since both the
volume and area are halved.
This bed depth is greater than the recommended maximum of 1.2 m (4 ft). This
depth would correspond to a pressure drop of approximately 10 kPa (40 in. H2O)
assuming the carbon to be 4 x 10 mesh from Figure 5-11. In practical design situa-
tions the operating parameters (such as bed depth, vessel size, cycle time, etc.)
would be varied to reduce this excessive pressure drop and to determine the most
cost effective alternatives.
The above solution is based on a number of adsorber design maximum and
minimum rules of thumb. It is intended as a guide to illustrate how to "red flag"
any parameters which may be greatly exceeded.
References
Bethea, R. M. 1978. Air Pollution Control Technology. New York: Van Nostrand
Reinhold.
Calgon Corporation. Air Purification with Activated Carbon, Technical Bulletin.
Pittsburgh.
Cannon, T. E. 1977. Carbon Adsorption Applications In Air Pollution Control and
Design Handbook, P. N. Cheremisinoff and R. A. Young, eds. New York:
Marcel Dekker, Inc.
5-38
-------
Cerny, S., and Smesek, M. 1970. Active Carbon. New York: Elsever Publishing
Company.
Chemical Engineering, 1977. Beaded Carbon Ups Solvent Recovery. August 29.
Dubinin, M. M. and Saverina, E. 1936. The Porosity and Sorptive Properties
of Active Carbon. Aota Physicochem. USSR, 4. 647.
Environmental Protection Agency (EPA). 1973. Package Sorption Device System
Study. EPA-R2-73-202. Research Triangle Park, NC.
Hellman, T. M. 1977. Odor Control By Adsorption. In Air Pollution Control and
Design Handbook, P. N. Cheremisinoff and R. A. Young, eds. New York:
Marcel Dekker, Inc.
Kovach, L. J. 1978. Gas-Phase Adsorption and Air Purification. In Carbon
Adsorption Handbook, P. N. Cheremisinoff and F. Ellerbush, eds. Ann Arbor:
Ann Arbor Science Publishers, Inc.
Langmuir, I. 1918. The Adsorption of Gases on the Plane Surfaces of Glass,
Mica and Platinum. J. Am. Chem. Soc. 40:1361.
Perry, J. H. ed. 1973. Chemical Engineers Handbook. 5th ed. New York: McGraw
Hill Book Co.
Parmele, C. S., O'Connell, W. L., and Basdekis, H. S. 1979. Vapor-Phase
Adsorption Cuts Pollution, Recovers Solvent. Chem. Engr. 86:58-70 (December
31, 1979).
Sutcliffe Speakman Co. 1963. Solvent Recovery with Active Carbon. Technical
Bulletin. Bronxville, New York.
Turk, A. 1977. Adsorption. In Air Pollution Vol. IV Engineering Control of Air
Pollution, A. C. Stern, ed. New York: Academic Press.
Union Carbide. Purasiv HR for Hydrocarbon Recovery. Technical Bulletin.
New York.
Vic Manufacturing. Carbon Adsorption/Emission Control Technical Bulletin.
Minneapolis.
Woods, F. J. and Johnson, J. E. 1964. NRL Report 6090.
5-39
-------
Chapter 6
Condensation
Introduction
Condensation is the process of reducing a gas or vapor to a liquid. Any gas can be
reduced to a liquid by lowering its temperature and/or increasing its pressure. The
most common approach is to reduce the temperature of the gas stream, since
increasing the pressure of a gas is very costly (EPA, 1973).
Condensers are simple, relatively inexpensive devices that normally use water or
air to cool and condense a vapor stream. Since these devices are usually not
capable of reaching low temperatures (below 80 °F), high removal efficiencies of
most gaseous pollutants are not obtained unless the vapors will condense at high
temperatures. Condensers are typically used as pretreatment devices. They are used
ahead of incinerators, absorbers, or adsorbers to reduce the total gas volume to be
treated by more expensive control equipment. Used in this manner, they help
reduce the overall cost of the control system.
Condensation Principles
When a hot vapor stream contacts a cooler surface, heat is transferred from the
hot gases to the cooler surface. As the temperature of the vapor stream is cooled,
the average kinetic energy of the gas molecules is reduced. Also, the volume that
these vapors occupy is reduced. Ultimately the gas molecules are slowed down and
crowded together so closely that the attractive forces (van de Waals" forces)
between the molecules cause them to condense to a liquid.
The two conditions which aid condensation are: low temperatures so that the
kinetic energy of the gas molecules are low; and high pressures so that the
molecules are brought close together. The actual conditions at which a particular
gas molecule will condense depends on its physical and chemical properties.
Condensation occurs when the partial pressure of the pollutant in the gas stream
equals its vapor pressure as a pure substance at operating conditions.
Condensation of a gas can occur in three ways: first, at a given temperature, the
system pressure is increased (compressing the gas volume) until the partial pressure
of the gas equals its vapor pressure; second, at a fixed pressure, the gas is cooled
until the partial pressure equals its vapor pressure; or third, by using a combina-
6-1
-------
tion of compression and cooling of the gas until its partial pressure equals its vapor
pressure. These processes are illustrated in Figure 6-1, a typical vapor pressure
diagram for a pure substance.
Critical
point
Cool
Temperature
Figure 6-1. Typical vapor pressure curve.
In Figure 6-1, point I is the initial temperature and pressure of a gas. The
dotted lines indicate the paths a quantity of gas would follow to reach the vapor
pressure curve. Points on the vapor pressure curve are also referred to as dew
points. The dew point is defined as the condition at which gas is ready to condense
into the first drop of liquid.
Also on Figure 6-1, the critical point is plotted. Each substance has a critical
temperature and critical pressure. The critical temperature is important in that it
is a maximum temperature above which the gas will not condense, no matter how
great a pressure is applied. The pressure required to liquify a gas at its critical
temperature is the critical pressure.
Once the gas conditions (temperature, pressure, and volume) equal those on the
vapor pressure line, liquid begins to condense. From this point on, the gas-liquid
mixture follows the vapor pressure line. If the mixture is cooled continuously, the
partial pressure of the remaining gas will always equal the vapor pressure. This is
6-2
-------
important since even though the contaminated gas is being condensed, it still hat. a
certain partial pressure indicating that uncontrolled vapors are being emitted from
the condenser. For most practical applications, the vapor-liquid equilibrium
restricts the use of condensers as primary air pollution control devices. Unless very
low temperatures or high pressures are attained, condensers are not capable of
reducing the pollutant concentration to within acceptable emission limits.
Practically, temperature is the only process variable which governs the effec-
tiveness of a condenser. In industrial applications, increasing the system pressure is
very costly and therefore rarely used for condensation. At the operating pressure of
the system, the outlet temperature from the condenser determines the maximum
removal efficiency. Therefore, condensers cannot be used in the same manner as
other gaseous pollutant control devices. For example, condensers cannot be used in
series like adsorbers or absorbers to further reduce outlet concentration unless the
outlet temperature of the second condenser is lower than the previous one.
Increasing gas residence time or decreasing flow rates in the condenser does not
add to the theoretical achievable efficiency as these operations do in incinerators,
adsorbers, and absorbers.
Condensers
Condensers fall into two basic categories; contact and surface condensers. In a con-
tact condenser the coolant and vapor stream are physically mixed. They leave the
condenser as a single exhaust stream. In a surface condenser, the coolant is
separated from the vapors by tubular heat transfer surfaces. The coolant and con-
densed vapors leave the device by separate exits. Surface condensers are commonly
called shell-and-tube heat exchangers. The temperature of the coolant is increased,
so these devices also act as heaters.
6-3
-------
Direct Contact Condensers
Contact condensers are simple devices such as spray towers steam or water jet
ejectors, and barometric condensers. These devices bring the coolant, usually
water, into direct contact with the vapors as illustrated in Figure 6-2. The liquid
stream leaving the condenser contains the coolant plus the condensed vapors. If the
vapor is soluble in the coolant then absorption also occurs. Absorption increases the
amount of contaminant that can be removed at the given conditions.
(a) spray contact condenser
Mist eliminator
Spray nozzles
Figure 6-2. Direct contact condensers:
(a) spray, (b) jet ejector, and (c) barometric.
6-4
-------
(b) jet ejector condenser
(c) barometric condenser
Spray
nozzle
Water
inlet
Water
inlet
Discharge
Discharge
Spray tower condensers (Figure 6-2a) are normally the same as the spray
absorbers discussed in Chapter 4. The vapors enter the bottom of the tower while
coolant is sprayed down over them. Baffles are usually added to ensure adequate
contact between coolant and vapors. Ejectors (Figure 6-2b) and barometric con-
densers (Figure 6-2c) operate in a similar manner. The difference being that they
use liquid sprays to move the vapor stream. In both of these devices the coolant is
sprayed into a venturi throat creating a vacuum which moves the vapor stream
through the condenser.
6-5
-------
Surface Condensers
Surface condensers are usually in the form of shell-and-tube heat exchangers
(Figure 6-3). The device consists of a circular or oval cylindrical shell into which
the vapor stream flows. Inside the shell are numerous small tubes through which
the coolant flows. Vapors contact the cool surface of the tubes, condense, and are
collected, while noncondensed vapors are sent for further treatment.
Removable
channel
cover
Reversing channel
Inlet
channel
Figure 6-3. Single-pass condenser.
Removable
channel
cover
Figure 6-3 is a diagram of a single-pass heat exchanger, where the entire stream
of coolant flows through all the parallel tubes. The cooling liquid enters and makes
one pass through the tubes or tube side of the exchanger. The uncondensed vapor
stream enters and makes one pass on the shell side of the exchanger. The single
pass exchanger is limited in that it requires a large number of tubes and low gas
velocities through the exchanger to provide adequate heat transfer. By using a
multipass system, shorter tube lengths, higher gas velocities through the exchanger,
and improved heat transfer can be achieved (McCabe, 1967).
6-6
-------
Figure 6-4 illustrates two types of multipass heat exchangers, referred to as 1-2
and 2-4 heat exchangers. The first digit refers to the number of passes the vapor
makes on the shell side, while the second digit indicates the number of tube side
passes. Both of these designs give improved performance over the single-pass
exchanger. The 2-4 heat exchanger is capable of higher gas velocities and better
heat transfer than the 1-2 heat exchanger. Adding more passes does have disadvan-
tages however. These disadvantages are: the exchanger construction is more com-
plicated; friction losses are increased due to the higher velocities; and exit and
entrance losses are multiplied.
Coolant Vapor
inlet inlet
Inlet
Coolant channel
outlet
1-2 parallel flow exchanger
Noncondensing
vapor outlet
Condensate
outlet
Reversing
channel
Coolant
inlet
Inlet
channel
Coolant
outlet
Vapor
inlet
2-4 exchanger
Straight
seamless tubes
FCondensate
outlet
Floating
head
Figure 6-4. Simplified air flow in multipass exchangers.
6-7
-------
Condensation applications normally require large temperature differentials
between vapor and coolant. These temperature variances can cause the tubes to
expand or contract. This expansion stress can eventually cause the tubes to buckle
or pull loose from the shell, destroying the condenser. Floating head construction is
commonly used to avoid condenser expansion stress damage. In a floating head,
one end of the tube bundle is mounted so that it is structurally independent from
the shell as shown in Figure 6-4. This allows the tubes to expand and contract
within the shell.
Water is generally the coolant used in condensers. However, short supply and
expense to treat water make it an uneconomical choice in some cases. In these
cases, air-coolers are used. The specific heat of air is only about 0.25 Btu/lb»°F,
approximately one-fourth that of water. Therefore, air condensers must be very
large compared to water condensers.
To conserve space and reduce the cost of equipment in these cases, heat
exchangers with extended surfaces can be used. In these devices, the outside area
of the tube is multiplied or extended by adding fins or disks. Figure 6-5 illustrates
two types of finned tubes. In extended surface condensers, the vapor is condensed
inside the tube while air flows around the outside contacting the extended surfaces.
Transverse fins
Longitudinal fins
Figure 6-5. Extended surface tubes.
Comparison of Contact and Surface Condensers
Since in contact condensers coolant is merely sprayed on the vapors, these systems
are simpler in design, less expensive, and more flexible in application than surface
condensers. However, contact condensers require more coolant, and due to direct
mixing, produce 10 to 20 times the amount of wastewater (condensate) than sur-
face condensers. Since the wastewater from a contact condenser is contaminated
with vapors, it cannot be reused posing a water disposal problem. If the condensed
vapors have a recovery value, surface condensers are usually used since the conden-
sate can be recovered directly.
6-8
-------
Design of Condensers
Condensers reduce the temperature (heat content) of a hot vapor stream by con-
tacting it with a cooler liquid or air stream. Heat is being transferred from a hot to
a colder fluid; thus the process is termed a heat transfer or heat exchange opera-
tion. Condensers are designed using the same basic principles and empirical rela-
tionships used for heat exchangers.
Procedures used to design simple heat exchangers are well defined and can be
found in numerous texts, such as: Perry (1973); McCabe (1967); or Kern (1950). In
adapting these procedures to designing a condenser, an additional variable com-
plicates the calculations. Heat exchanger designs are based on heat transfer theory.
In a condenser, however, both heat and mass are being transferred. As the vapor
stream passes through the condenser, both its temperature and its composition are
changing. Even the simplest of condenser design procedures found in Perry (1973)
or Bell (1972) are complicated and require trial and error solutions. A rigorous
procedure for the design of condensers will not be presented in this chapter.
Simplified heat balance calculations are used to estimate the important parameters.
The following equations are not intended to be used as a design method, but only
as rough estimates for evaluation purposes.
Heat Balance
The first step in analyzing any heat transfer process is to set up a heat balance
relationship. For a condensation system, the heat balance can be expressed as:
Heat in = Heat out
Heat required to reduce Heat required Heat needed to be
vapors to the dew point + to condense = removed by the coolant
vapors
This heat balance is written in equation form as:
(Eq. 6-1) q = rhCp(TG1 - TdeM point) + rhHv = LCP(T£2 - Ttl)
Where: q = heat transfer rate, Btu/hr*
rh = mass flow rate of vapor, Ib/hr
L = mass flow rate of liquid coolant, Ib/hr
Cp = average specific heat of a gas or liquid, Btu/lb»°F
T = temperature of the streams; G for gas and L for liquid coolant, °F
Hv = heat of condensation or vaporization, Btu/lb
*The units used in this chapter are in English, since most readily available heat transfer data is in
English units.
6-9
-------
In Equation 6-1, the mass flow rate (rh) and inlet temperature (TC1) of the vapor
stream are set by the process exhaust stream. The temperature of the coolant
entering the condenser (TZ1) is also set. The average specific heats (Cp) of both
streams, the heat of condensation (Hv), and the dew point temperature can be
obtained from chemistry handbooks. Therefore, only the amount of coolant (L)
and its outlet temperature are left to be determined. If either one of these terms
are set by process restrictions (i.e. only x pounds an hour of coolant are available
or the outlet temperature of coolant must be below a set temperature), then the
other term can be solved for directly.
Equation 6-1 is applicable for direct contact condensers and should be used only
to obtain rough estimates. Equation 6-1 has a number of limitations: the specific
heat (Cp) of a substance is dependent on temperature; and the temperature
throughout the condenser is constantly changing. Also, the dew point of a
substance is dependent on its concentration in the gas phase; and since rh is con-
stantly changing (vapors being condensed) the dew point temperature is constantly
changing. Finally, no provision is made for cooling the vapors below their dew
point. An additional term would have to be added to the left side of Equation 6-1
to account for this amount of cooling.
Surface Condensers
In a surface condenser or heat exchanger, heat is transferred from the vapor
stream to the coolant through a heat exchange surface. The rate of heat transfer
depends upon three factors: total cooling surface available, resistance to heat
transfer, and mean temperature difference between condensing vapor and coolant.
This can be expressed mathematically by:
(Eq. 6-2) q=UAATm
Where: U = overall heat transfer coefficient, Btu/°F«ftJ»hr
A = heat transfer surface area, ft2
ATm = mean temperature difference, °F
6-10
-------
U, the overall heat transfer coefficient is a measure of the total resistance heat
experiences while being transferred from a hot body to a cold body. In a shell-and-
tube condenser, cold water flows through the tubes causing vapor to condense on
the outside surface of the tube wall. Heat is transferred from the vapor to the
coolant. The ideal situation for heat transfer would be where heat is transferred
from the vapor to the coolant without any heat loss (heat resistance). Figure 6-6
shows the typical heat resistances occurring during heat transfer in a shell-and-tube
condenser. Every time heat moves through a different medium, it encounters a dif-
ferent and additional heat resistance. These heat resistances (Figure 6-6) are:
throughout the condensate, through any scale or dirt on the outside of the tube
(fouling), through the tube itself, and through the film on the inside of the tube
(fouling). Each of these resistances are individual heat transfer coefficients and
must be added together to obtain an overall heat transfer coefficient.
Cooling liquid
Inside fouling (scale or dirt)
Tube wall
Outside fouling (scale or dirt)
Layer of condensate
Vapor
Figure 6-6. Resistance to heat transfer around a cooling tube.
6-11
-------
A number of correlations exist to determine the individual heat transfer coeffi-
cient (Perry, 1973 and Kern, 1950). These correlations are usually presented in
terms of dimensionless numbers such as the Reynolds Number. These correlations
consider variables including geometry of the condenser, gas and liquid density and
viscosity, and the mechanism of heat transfer (convection, conduction and radia-
tion). Calculation of individual heat transfer coefficients require numerous data
from the system. An estimate of an overall heat transfer coefficient can be used for
preliminary calculations. Table 6-1 lists some typical values for overall heat transfer
coefficients. Table 6-1 should be used only for preliminary estimating purposes.
Table 6-1. Typical overall heat transfer coefficients in tubular heat exchangers.
Condensing vapor*
(shell side)
Alcohol vapor
High boiling hydrocarbons
(vacuum)
Low boiling hydrocarbons
Organic solvents
Organic solvents with high
percent of
noncondensables present
Naptha
Stabilizer reflux vapors
Sulfur dioxide
Tall oil derivatives,
vegetable oil vapors
Steam
Cooling liquid
(tube side)
Water
Water
Water
Water
Water or brine
Water
Water
Water
Water
Feedwater
U
(Btu/°F.ft'.hr)
100-200
20-50
80-200
100-200
20-60
50-75
80-120
150-200
20-50
400-1000
*Note: for water-water (liquid-liquid) heat exchanger (no phase change) the values for U range
between 200-250.
In a surface heat exchanger the temperature difference between the hot vapor
and the coolant usually varies throughout the length of the exchanger. Therefore,
a mean temperature difference (ATm) must be used. For the special cases where the
flow of both streams is completely cocurrerit, the flow of both streams is com-
pletely countercurrent, or the temperature of one of the fluids remains constant (as
is the case in condensing a pure liquid), the log mean temperature difference can
be used. The temperature profiles for these three conditions are illustrated in
Figure 6-7. The log mean temperature for countercurrent flow can be expressed as:
(Eq. 6-3)
Where: AT,m = log mean temperature
6-12
-------
(a) cocurrent flow
3
I
H
(b) countercurrent flow
Ll
(c) isothermal condensation with countercurrent flow
C1
dew point
Length of exchanger
Figure 6-7. Temperature profiles in a heat exchanger for countercurrent flow,
cocurrent flow, and isothermal condensation with countercurrent flow.
6-13
-------
The value calculated from Equation 6-3 is used for single pass heat exchangers or
condensers. For multiple pass exchangers a correction factor to the log mean
temperature must be included. However, for the special case of isothermal conden-
sation (no change in temperature) of a single component vapor, T0i = Tcz and is
equal to the dew point temperature. No correction factor is needed (Perry, 1973).
In order to size a condenser, Equation 6-2 must be rearranged to solve for the
surface area.
(Eq. 6-4) A=
UAT,m
Where: A = surface area of a shell-and-tube condenser, ftz
q = heat transfer rate, Btu/lb
U = overall heat transfer coefficient, Btu/°F«ftz«hr
AT,m = log mean temperature, °F
For the exact determination of the surface area, U in Equation 6-4 should be
calculated from the individual resistances (coefficients). For a rough estimate the
values in Table 6-1 can be used.
Equation 6-4 is valid only for isothermal condensation of a single component.
This implies that the pollutant is a pure vapor stream comprised only of one
specific hydrocarbon such as benzene, not a mixture of hydrocarbons. Nearly all
air pollution applications involve multicomponents since air and any other gas are
two-components. This complicates the design procedure for a condenser.
If the incoming vapor stream is above its dew point (superheated), it must first
be cooled to its dew point. It is generally considered conservative to size the con-
denser as if the entire heat load (desuperheating and condensing) were transferred
by condensation (Perry, 1973). In this case, the heat load, q, is computed by using
the left side of Equation 6-1. The dew point temperature of the gas is used in
calculating the log mean temperature.
Solving for the surface area, A, this method will give the size of the condenser
that will cool the vapor to the dew point and condense it, although this method
will not include the heat removal required for subcooling. The liquid must be sub-
cooled to ensure that vapors will not readily volatilize from the condensate. Sub-
cooling the liquid is accomplished by maintaining a liquid level which covers some
of the bottom tubes in the condenser. This condition is called flooding.
The single phase (liquid) subcooling section can be treated separately, giving an
area that is added onto the area needed for condensation. The heat balance and
surface area are determined on the basis of cooling the liquid from its dew point
temperature to the subcooled outlet temperature.
Different situations can occur when trying to condense vapors containing more
than one component. In one case, all the components will condense at the coolant
temperature. In another case, a few of the components may condense, while the
other components will not condense as pure components but be very soluble in the
condensate. Finally, there are mixtures where one or more of the components will
not condense and are insoluble in the liquid condensate.
6-14
-------
No matter which of these situations occur, features common to all three can
affect condenser design. Features which must be considered are: multicomponent
condensation is always nonisothermal; both liquid and vapor sensible heat removal
must be provided for; and the composition of both phases is continually changing
(Bell, 1972). This last feature is what complicates the design of multicomponent
condensers. In multicomponent condensers, the liquid and gas phases are con-
tinually changing. This results in diffusional resistances occuring in both the liquid
and vapor phases in addition to the resistance to heat transfer. In order to design
the condenser, temperature and concentration profiles and diffusion coefficients
must be evaluated at various points in the condenser. Methods for designing
multicomponent condensers will not be discussed here, but are presented in Perry
(1973) and in Bell (1972).
In nearly all condensation applications related to air pollution control, the gas to
be treated is a mixture of condensable and noncondensable vapors. Nonconden-
sable gases offer resistance to the rate of condensation in terms of both heat and
mass transfer since the condensable vapor must diffuse to the cool surface. Air is
usually the contaminant vapor but other gases that do not condense or dissolve in
the liquid will also affect heat transfer. Unless specifically accounted for, the
presence of noncondensable fluids will lower the values of the overall heat transfer
coefficients that are listed in Table 6-1.
For preliminary rough estimates of condenser size, the procedure in Example 6-1
for single component condensation can be used for condensing multicomponents.
In choosing a heat transfer coefficient, the smallest value should be used to allow
as much overdesign as possible.
Example 6-1
In a rendering plant, tallow is obtained by removing the moisture from animal
matter in a cooker. Exhaust gases from the cookers contain essentially steam,
however the entrained vapors are highly odorous and must be controlled. Con-
densers are normally used to remove most of the moisture prior to incineration,
scrubbing, or carbon adsorption.
The exhaust flow rate from the continuous rendering cooker is 20,000 acfm at
250 °F. The exhaust gases are 95% moisture with the remaining portion consisting
of air and obnoxious organic vapors. The exhaust stream is sent first to a shell-and-
tube condenser to remove the moisture and then to a carbon adsorption unit. If
the coolant water enters at 60 °F and leaves at 120°F, estimate the required surface
area of the condenser. The condenser is a horizontal, countercurrent flow system
with the bottom few tubes flooded to provide subcooling.
6-15
-------
Solution:
1. First, compute the pounds of steam condensed per minute.
20,000 acfmx 0.95= 19,000 acfm steam
I „ ,
From the ideal gas law:
PV = nRT
PV _ (1 atm)(19.000 acfm)
RT
/Q atm.fta \
\ lb mole'OR/
= 36.66 Ib mole/min
• Ua cc lb rnoleV 18 lb \
m = (36.66 - : — I -
\ min /\lb mole/
= 660 Ib/min of steam to be condensed
2. Solve the heat balance to determine q for cooling the superheated steam and
condensing only.
heat needed to
q= cool steam to + heat of
condensation temperature condensation
= rhCpAT + mHv
The average specific heat (Cp) of steam @ 250°F«0.45 Btu/lb»°F. The
heat of vaporization of steam @ 212 °F= 970.3 Btu/lb. (Both these values
were obtained from Perry, 1973).
Substituting into the equation:
q= (660^(0.45-^) (250-212°F)+(660 ^
\ mm/V lb°F/ \ mi
.
min/\ lb /
= 11,286 Btu/min + 640,398 Btu/min
= 651,700 Btu/min
3. Now using Equation 6-4 to estimate surface area for this part of the condenser:
A= H
For a countercurrent condenser the log mean temperature is given by:
AT = (Tci ~ TLZ) — (TC2 — T£1)
In Tci ~
TGI ~
6-16
-------
Remember that the desuperheater-condenser section is designed using the
saturation temperature in calculating the log mean temperature difference.
TG, = 212°F
3 Ti2=120°F
a
I
4J
h
, = 212°F
Length of exchanger
AT = (212-120)-(212-60)
212-120
In
212-60
= 119.5°F
The overall heat transfer coefficient, U is assumed to be 100 Btu/°F»ft*»hr
from Table 6-1 for condensing steam. This value is chosen from the line in the
table for stabilizer reflux vapors. Stabilizer reflux vapors are a mixture of
hydrocarbon and steam similar to that exhausted from a rendering dryer. The
median value of 100 is used as a best guess to approximate the conditions.
Substituting the appropriate values into Equation 6-4:
/60 min\
A= (651.700 Btu/min) \ hr
(100 Btu/0F.ft2-hr)(119.5°F)
= 3272 ft2
4. To estimate the total size of the condenser, we need to allow for subcooling
of the water (212°F-160°F). (160°F is a safe margin).
Refiguring the heat balance for cooling the water:
q = mCpAT
Where: rh * 660 Ib/min (assuming all the steam is condensed)
Cp for water = 1 Btu/lb»°F
= 34,320 Btu/min
6-17
-------
The new log mean temperature is:
TC1 = 212°F
3 TL2=120°F
U
SL
I
TC2=160°F
T£1 = 60°F
Length of exchanger
AT = (212-12Q)-Q60-60)
212-120
160-60
= 96°F
A=
UAT/m
U for cooling water with a water coolant is 200 Btu/°F«ftz«hr from the footnote
at the bottom of Table 6-1.
A =
34,320 Btu/min
(200 Btu/°F.ft2«hr)(96°F)
= 1.79 ft2 or 2 ft2
5. The total area needed is:
A = 3272+ 2
= 3274 ft2
As illustrated by this example, the area for subcooling is usually very small com-
pared to the area required for condensing.
References
Bell, K. J. and Ghaly, M. A. 1972. An approximate generalized design method
for multicomponent/partial condensers. Am. Inst. Chem. Engrs. Symp. Series.
Vol. 69. No. 131, pp. 72-79.
Kern, D. Q. 1950. Process Heat Transfer. New York: McGraw Hill Book Co.
6-18
-------
Lord, R. C., Minton, P. E., Silusser, R. P. 1970. Design of Heat Exchangers.
Chem. Engr. 77:96-118 (Jan. 26).
Marchello, J. M. 1976. Control of Air Pollution Sources. New York: Marcel
Dekker, Inc.
McCabe, W. L. and Smith, C. J. 1967. Unit Operations of Chemical Engineering.
New York: McGraw Hill Book Co.
Perry, J. H. ed. 1973. Chemical Engineers Handbook 5th ed. New York: McGraw
Hill Book Co.
Peters, M. S., and Timmerhaus, K. D. 1968. Plant Design and Economics for
Chemical Engineers. New York: McGraw Hill Book Co.
Standiford, F. C. 1979. Effect of Non-Condensables on Condenser Design and
Heat Transfer. Chemical Engineering Progress, pp. 59-62.
6-19
-------
Chapter 7
Control of Nitrogen Oxide Emissions
Introduction
More than 20 million metric tons of nitrogen oxides (NOX) are emitted into the
atmosphere each year as a result of burning fossil fuels (EPA, 1980). Fires and
other natural occurrences emit 10 times the amount that anthropogenic nitrogen
oxides emit, and tend to be dispersed over very large areas. Nitrogen oxides emit-
ted from combustion processes can create local ambient levels that are 10 to 100
times greater than natural concentrations (EPA, August 1980). Controlling
nitrogen oxides is important because of their harmful effects on public health and
welfare, and their involvement in the formation of photochemical smog.
Sources of Nitrogen Oxides
Approximately one half of the total nitrogen oxide emissions come from motor
vehicles. The other half comes from fossil fuel combustion by utility and industrial
boilers and other industrial furnaces and processes (Figure 7-1). EPA has estimated
that stationary source nitrogen oxide emissions will grow to 70% of the total by
1985 (EPA, August 1980). Part of this growth is due to the increased use of coal
for generating electricity. There is more nitrogen in coal than in most other fossil
fuels. Nitrogen oxides are released as coal is burned.
Of the stationary sources producing nitrogen oxides, utility and industrial boilers
are the most significant; emitting approximately 60% of the total emissions (EPA,
January 1978).' Only about 2% of the NO, emissions come from industrial processes
such as nitric acid plants. The remaining stationary source NO, emissions are from
industrial, commercial, and residential heating. Emissions from nitric acid plants
are usually reduced by control methods such as wet scrubbing or catalytic reduc-
tion. NO, emissions from combustion sources are reduced by combustion modifica-
tions, flue gas treatment, or fuel denitrification. These reduction techniques will be
discussed later in this chapter.
7-1
-------
Electric generation (33%)
Industrial (12%)
Residential
1%
Commercial and
institutional
(4%)
Source: EPA, February 1980
Figure 7-1. Manmade nitrogen oxide emissions.
Effects
NO, emissions are known to cause human health problems and environmental
damage. Exposure to nitrogen oxides is believed to increase the risks of acute
respiratory disease and susceptibility to chronic respiratory infection. Nitrogen
dioxide (NOZ) contributes to heart, lung, liver and kidney damage. Exposure to
high NO2 concentrations can be fatal. At lower NO2 levels, 25 to 100 parts per
million, acute bronchitis and pneumonia can occur. Eye and skin irritations can
result from occasional exposure to low concentrations of NOj, (EPA, August 1980).
Nitrogen oxides are also toxic to vegetation. Reduced growth and plant destruction
have been observed as a result of exposure to NO, emissions.
Probably the most important effect of NO, is the contribution to photochemical
smog. NO, reacts with hydrocarbons in the presence of sunlight to form
photochemical oxidants, or ozone (Os). Ozone can cause severe damage to the
lungs and can aggravate asthma and other respiratory diseases. In addition to
ozone, other smog pollutants such as peroxyacetylnitrate (PAN) can cause
coughing, eye irritation, headaches, and throat irritation.
Nitrogen oxides also contribute to the global acid rain problem. Through a
series of complex reactions, nitrogen oxides can be converted to nitric acid and
nitrates which may be deposited in the form of rain and snow. EPA has estimated
that nearly one-half of the acidity in acid rain is due to nitric acid (EPA, August
1980).
7-2
-------
New Source Performance Standards
EPA has promulgated New Source Performance Standards (NSPS) lor fossil-fuel-
fired steam generators (FFFSG) with heat input greater than 250 X 10s Btu/hr
(73 MW thermal). These standards establish nitrogen oxide emission limits for
utility boilers. There are two standards: one for steam generators installed after
August 17, 1971; and one for steam generators installed after September 18, 1978
(Table 7-1). These standards were set to reduce the NOX emissions from the largest
combustion source, utility boilers. Standards for industrial boilers are forthcoming
in the near future.
Table 7-1. New Source Performance Standards for NO, emissions from
fossil-fuel-fired steam generators rated at greater than
250 x 10" Btu/hr or 73 MW (thermal) heat input.
NO, emissions
Subpart D; new
sources after
August 17, 1971
Subpart Da; new
sources built after
September 18, 1978
Metric units
(ng/J)
86
130
300
86
130
210
260
340
English units
(Ib/lO* Btu)
0.2
0.3
0.7
0.2
0.3
0.5
0.6
0.8
Fuel
Gaseous
Liquid
Solid (except lignite)
Gaseous (except coal derived)
Liquid (except coal derived)
Subbituminous coal
Bituminous/ anthracite
coal and lignite
Lignite mined in ND, SD,
and Montana burned in
a slag-top furnace
Research Activities
The EPA's Office of Research and Development (ORD) is currently developing
inexpensive methods to reduce NO* emissions from combustion sources. Most recent
studies for NO* control technology are being done by the EPA's Industrial
Environmental Research Laboratory in Research Triangle Park, NC (IERL-RTP).
Additional research is being conducted at the Environmental Sciences Research
Laboratory in Research Triangle Park, NC (ESRL-RTP) and through the Head-
quarters Office of Environmental Engineering and Technology in Washington, DC
(EPA, August 1980).
The most promising control methods for reducing NO, emissions from combus-
tion sources are:
Before burning:
• fuel denitrogenation (denitrification)
During burning:
• combustion modifications
• catalytic combustion
After burning:
• flue gas treatment
• catalytic emission control
7-3
-------
Fuel Denitrogenation
One control technology for reducing NO* emissions is to remove the nitrogen con-
tained in the fuel. Amounts of nitrogen vary in a fossil fuel. Coal, shale, and
residual fuel oil contain a larger amount of nitrogen than either distillate oil or
natural gas. The nitrogen in the fuel can be emitted as NO* when the fossil fuel is
burned in the furnace.
Nitrogen is removed from coal, shale, or heavy fuel oils by liquifying the fuels
and mixing with hydrogen gas. The mixture is heated and a catalyst is used to
cause the nitrogen in the fuel and the hydrogen to unite. This reaction produces
two products; ammonia and a cleaner fuel. Researchers are developing better
catalysts and finding ways to reduce the deposition of carbon on the catalyst sur-
face (EPA, August 1980). Carbon deposits reduce the effectiveness of the catalyst
life (see Chapter 3). This technology can reduce the nitrogen content in both
natural fuels and synthetic liquid or gaseous fuels (made from shale and coal). This
could become an increasingly important technology with the development and use
of synthetic fuels in the future.
Formation of Nitrogen Oxides
in Combustion Sources
When fossil fuels are burned with air in a furnace, some of the oxygen (O2) and
nitrogen (N2) present combine to form oxides of nitrogen. Most of the oxides form
according to the following reaction:
Once the NO forms, the rate of decomposition is very slow and NO does not
dissociate into N2 and O2 in any appreciable amounts.
The NO formed can react with more oxygen to form NOZ.
In large combustion furnaces, the majority of NO* formed, approximately 95%, is
in the form of NO. The main factors involved in NO* formation are: the flame
temperature; the length of time combustion gases are maintained at that
temperature; and the amount of excess air present in the flame. Flame
temperature in a utility boiler furnace is approximately 1650°C (SOOO°F). At this
high temperature NO is formed in great abundance (sometimes greater than
1000 ppm). The residence time usually available in combustion equipment is too
short for an appreciable fraction (usually less than 5%) of the NO to be oxidized to
NO2. The bulk of the NO2 is formed in the atmosphere after release from the stack
rather than in the confines of the combustion furnace. In this manual, NO* refers
to the total NO and NO2 being emitted from major combustion sources.
7-4
-------
In small residential heating units, furnace temperatures are usually less than
1090°C (2000 °F). These relatively low temperature-operating units produce an
exhaust gas containing low NO, emissions; usually less than 10 ppm (Public Health
Service, 1970).
Thermal NOX and Fuel NOX
Nitrogen oxide emissions are formed by two chemical processes occurring during
combustion: the thermal NO, process and the fuel NO, process.
Thermal NO, results from intense heat during combustion causing the nitrogen
content of the combustion air to be oxidized. Its rate of formation is highly sen-
sitive to the flame temperature and to a lesser extent to the local concentration of
oxygen at the flame. Virtually all thermal NO, is formed in the region of the flame
which is at the highest temperature (EPA, January 1978). In theory, the formation
of thermal NO* can be reduced in four ways: reduce the nitrogen level (of air) at
peak temperature; reduce the oxygen level at peak temperature; reduce the
peak flame temperature; and reduce the time of exposure at peak temperature
(EPA, January 1978). Since the concentration of nitrogen in air-fuel mixtures is
relatively fixed, this tactic is not really applicable. Therefore, thermal NO, is
reduced in field practice by reducing oxygen levels, peak flame temperatures, and
residence time in the NO, producing section of the furnace. This is accomplished
by various combustion modification techniques such as the use of low excess air,
staged combustion, reduced air preheat, or flue gas recirculation.
Fuel NO, occurs as the nitrogen contained in the fuel is oxidized. Therefore, it is
dependent on the nitrogen content of the fuel. In fuels such as coal and heavy oil
that are high in nitrogen content, approximately 20 to 60% of the fuel-bound
nitrogen is oxidized (EPA, January 1978). Its rate of formation is strongly affected
by the local oxygen concentration present in the flame and also by the mixing rate
of the fuel and air. Thus, like thermal NO,, fuel NO, is dominated by the local
combustion conditions. One way to reduce fuel NO, emissions is to reduce the
nitrogen content in the fuel. This is not always possible and therefore combustion
modification techniques are used to reduce NO, emissions. These include com-
bining the use of low excess air firing, optimum burner designs, two stage combus-
tion or high air preheat (secondary air preheat).
Combustion Modifications
As previously discussed, nitrogen oxide emissions result from operating conditions
in the furnace, the amount of nitrogen in the combustion air, and the amount of
nitrogen in the fuel. Combustion conditions in the furnace can be modified to
reduce NO, emissions. Some of the more widely used combustion modification
techniques will be discussed in this section.
7-5
-------
A typical furnace with boiler tube sections is shown in Figure 7-2. The main sec-
tions are the fire wall or water wall tubes, superheater, convection tubes,
economizer, and air preheater. These boiler sections transfer heat in the furnace to
produce steam for generation of electricity or process steam in an industrial plant.
The furnace is surrounded by water walls or fire wall tubes. Water fed to the tubes
is heated and is turned into steam. Steam is collected in steam drums located in
the main convection section of the furnace. Convection tubes are used to convert
any water collected in the steam drum back into steam. The superheater is a sec-
tion of tubes used to superheat the steam to a higher temperature prior to going to
the turbine. The economizer is used to heat the feedwater makeup to the boiler.
The air preheater is used to preheat the air used for combustion. The furnace is
designed to use as much of the heat as possible before exhausting flue gas out the
stack.
Superheater
Convection
tubes
Steam
drum
Economizer
Air preheater
Figure 7-2. Typical boiler.
7-6
-------
Low Excess Air
In a combustion system, a certain amount of excess air is required to ensure com-
plete combustion of the fuel (Chapter 3). The more efficient the burners are for air
and fuel mixing, the less amount of excess air required for complete combustion
(Figure 7-3). The minimum amount of excess air is limited by the production of
smoke and unburned fuel leaving the furnace.
Normal or high
excess air
Low excess air
Figure 7-3. Excess air.
The level of excess air in an industrial or utility boiler will usually range from
5% to as high as 50 or 100%. However, most large boilers today operate with
excess air less than 5%. Small residential heating furnaces, because of their
unsophisticated design, usually operate with approximately 80 to 100% excess air.
Gas turbines operate at very high excess air conditions, 300 to 400%. Most of this
excess air is added as secondary air at lower temperatures resulting in low NOX
emissions.
NO, emissions are reduced from many combustion furnaces by low excess air
firing. The local flame zone concentration of oxygen is reduced, thus reducing
both thermal and fuel NO*. This method is easy to implement and actually
increases the efficiency of the furnace slightly. However, there are problems with
this combustion modification. Low excess air firing can produce smoke and high
CO emissions. Also, fouling and slagging of boiler tube surfaces can occur if
various coals and residual oils are burned.
NO, reductions averaging between 16 and 20% are achieved on gas and oil fired
utility boilers when the excess air is reduced to a level between 2 and 7% (EPA,
January 1978). NO* reductions averaging around 20% can be achieved on coal
fired utility boilers if the excess air is reduced to the 20% level or lower (EPA,
January 1978). For most utility boiler applications, operating at low excess air con-
ditions is considered a routine procedure.
7-7
-------
Staged Combustion
During staged combustion, air and fuel mixtures are combusted in two separate
zones. In one zone, the fuel is fired with less than a stoichiometric amount of air.
This creates a fuel rich local zone in the regions of the primary flame. The second
zone is an air rich zone where the remainder of the combustion air is introduced
to complete the combustion of the fuel (Figure 7-4). The heat in the primary flame
zone is not as intense as with normal firing because combustion is incomplete. The
air mixed with the fuel is sub-stoichiometric in the NO* forming region of the
flame, thus creating a low NO* condition. This modification is also referred to as
off-stoichiometric combustion.
Fuel rich
zone
Air rich
zone
Figure 7-4. Staged combustion.
Staged combustion reduces NO* emissions by a combination of several factors.
First, a lack of available oxygen for NO* formation in the fuel rich stage is due to
off-stoichiometric firing. Second, the flame temperature may be lower in the first
stage than with single stage combustion. Third, the peak temperature in the second
stage (air rich) is lower.
Staged combustion is an effective technique for controlling both thermal and
fuel NO* due to its ability to control the mixing of fuel with combustion air. The
NO* reduction effectiveness depends on good burner operation to prevent convec-
tive boiler tube fouling, unburned hydrocarbon emissions, and poor ignition
characteristics which occasionally occur at an excessively fuel rich boiler operation.
Fire side boiler tube corrosion may occur when burning some coals or heavy oils
under staged combustion conditions (EPA, September 1975).
7-8
-------
In staged combustion the flame is long, yellow, and slightly smokey as opposed
to the short and intense flame observed on normal firing. Fuel combustion (second
stage) also extends further into the furnace, sometimes causing excessive
temperatures in the convective and superheater sections of the boiler.
Staged combustion is accomplished by using modifications to the boiler such as
over fire air ports (OFA), or burners out of service (BOOS). Over fire air ports are
separate ports located above the burner. The burners are operated fuel rich with
the remainder of the combustion air coming in through the over fire air ports
(Figure 7-5). These are sometimes referred to as NOX ports. In some boilers, a
number of the burners are operated fuel rich, others air rich in a staggered con-
figuration. This is called biased firing. In the case where some burners are
operated on air only, this modification is called burners out of service (Figure 7-6).
NO* ports
Figure 7-5. Over fire air.
7-9
-------
Figure 7-6. Burners out of service.
On existing boilers, a steam load reduction will result with burners out of service
if the active fuel burners do not have the capacity to supply fuel for a full load.
Most utility boilers installed since 1971 have been designed with over fire air ports
so that all fuel burners are active during the staged combustion operation (EPA,
January 1978). Using staged combustion modifications on oil and gas fired boilers
reduces NO* emissions by approximately 30 to 40% (EPA, 1978). Modifying
existing coal boilers has reduced NO* emissions 30 to 50% (EPA, August 1980).
Staged combustion can also be accomplished by careful control of air and fuel
mixing in the burner. EPA (IERL) is working on the development of new coal
burners based on a staged combustion design that may reduce NO* as much as 85%
(EPA, August 1980). These low NOX burners will be discussed in more detail later
in this chapter.
Flue Gas Recirculation
Flue gas recirculation (FGR) has been used to reduce thermal NO* emissions from
large oil and gas fired boilers. A portion (10 to 30%) of the flue gas exhaust is
recycled back into the main combustion chamber by removing it from the stack
breeching and mixing it with the secondary air windbox (Figure 7-7). In order for
FGR to be effective in reducing NO, emissions, the gas must enter directly into the
combustion zone. This recirculated gas lowers the flame temperature and dilutes
the oxygen content of the combustion air, thus lowering thermal NO* emissions.
7-10
-------
Secondary
air wind
box
Stack
breeching
Figure 7-7. Flue gas recirculation.
Some operational problems can occur using flue gas recirculation. Possible flame
instability, loss of heat exchanger efficiency; and for small package boilers, conden-
sation on internal heat transfer surfaces limit the usefulness of gas recirculation
(EPA, January 1978).
Flue gas recirculation requires greater capital expenditures than low excess air
and staged combustion modifications. High temperature fans (forced or induced
draft), ducts, and large space are required for recirculating the gas. However, the
costs can be reasonable if the boiler is already equipped with gas recirculation for
steam temperature control (EPA, September 1974).
NOX reduction of approximately 40 to 50% is possible with recirculation of 20-to
30% of the exhaust gas in gas and oil fired boilers. Since FOR is used to reduce
thermal NO*, it is not effective for reducing NO* emissions generated from burning
a high nitrogen content fuel such as coal. At high rates of recirculation (>30%)
the flame can become unstable, thus increasing carbon monoxide and hydrocarbon
emissions.
The combustion modification technique (FGR) should not be confused with flue
gas recirculation used by utilities to control boiler tube (superheat) temperature. In
this utility practice, flue gas is taken from the stack breeching and mixed into the
hopper bottom section of the boiler that is located below the combustion chamber.
Since flue gas is not recirculated into the main combustion chamber, this utility
practice is not useful in reducing NO* emissions.
7-11
-------
Low NOX Burners
New low NO, burners have been developed by several manufacturers to reduce NO*
emissions. Burners control mixing of fuel and air in a pattern to keep flame
temperature low and dissipate the heat quickly. Some burners are designed to con-
trol the flame shape to minimize the reaction of nitrogen and oxygen at peak flame
temperatures. Other designs have fuel rich and air rich regions to reduce flame
temperature and oxygen availability.
The Dual Registered Low NO* burner is shown in Figure 7-8. The mixture of
pulverized coal and primary air is controlled to slightly delay combustion of the
fuel. The remainder of the combustion air (secondary air) is introduced through
two concentric air zones which surround the coal nozzle. The flame produced is
elongated compared to the short intense flame produced in a conventional burner.
The peak flame temperature is reduced, thus lowering thermal NO* emissions; and
the oxygen available in the flame is lower, thus reducing fuel NO, emissions. Field
tests of this burner have shown a 40 to 60% reduction in NO* emissions
(EPRI, 1979).
Secondary air
Pulverized coal
and primary air
Figure 7-8. Schematic of a low NO* burner.
EPA is developing a low NO, burner that produces a fuel rich primary combus-
tion zone and controls air and fuel mixing. Additional air is introduced from the
periphery of the burner to complete combustion in a secondary zone. These condi-
tions lead to a reducing atmosphere and a preferential conversion of the nitrogen
in the coal to molecular nitrogen (N2) emissions. This burner design may reduce
NOX emissions as much as 85% (EPA, August 1980).
7-12
-------
Other Combustion Modification Techniques
Other combustion modifications can also reduce NO, emissions from combustion
sources. Some of these include reduced air preheat, load reduction, steam and
water injection, and catalytic combustion.
Reduced air preheat and load reduction are used sparingly in large boilers due
to the energy penalty involved and the relatively low emission reduction occurring.
Steam and water injection are used mainly for NO, emission reduction in gas tur-
bines and internal combustion engines. Steam or water is injected into the combus-
tion area to lower the peak flame temperature and thus reduce thermal NO, emis-
sions. In catalytic combustion, a catalyst is used to achieve oxidation of the fuel
rather than using high flame temperatures. These systems have been used in gas
turbines to reduce NO, emissions well below 10 ppm (EPA, August 1980).
Flue Gas Treatment
Nitrogen oxide emissions can be reduced by treating the flue gas after it leaves the
combustion zone. Flue gas treatment processes include Exxon Thermal DeNOx,
Selective Catalytic Reduction (SCR), Shett UOP, and Wet Simultaneous NO* and
SO2 reduction. These processes have been used in Japan to reduce NO, emissions
from utility boilers and pilot projects are currently being tested in the United States
at various utilities. Full scale units are expected to be installed in the next few
years.
Exxon Thermal DeNOx Process
A process using ammonia injection to reduce NO, emissions has been developed by
Exxon Research and Engineering Co. called Thermal DeNO,. Ammonia is injected
into the post combustion zone of the boiler (Figure 7-9). The ammonia reacts with
NO, (which is 95% NO) to reduce the oxides to molecular nitrogen and water:
(Eq- 7-1) 4NH, + 4NO + 02-4N2 + 6H20
The above reaction is extremely temperature dependent. In a boiler, this reaction
successfully takes place at approximately 950 °C (1740°F). At higher temperatures
(above 1090°C) the ammonia is oxidized, forming additional NO,. At lower
temperatures (below 850 °C) the ammonia passes through the boiler unreacted
(Lyon and Tenner, 1978). This temperature range can be altered somewhat with
the addition of hydrogen. Hydrogen shifts the lower temperature range where the
reaction occurs at approximately 700°C. The hydrogen does not widen the range,
but just changes it. At H2:NH3 ratios of 2:1, the reaction will proceed rapidly at
700°C (Lyon and Tenner, 1978). By selecting various H2:NH3 ratios, the NO,
reduction can be accomplished at intermediate temperatures.
7-13
-------
This process also depends on the ammonia and NO* concentrations in the
system. As can be seen from Equation 7-1, four moles of NHS are needed to reduce
4 moles of NO, or, a 1:1 mole ratio. Injection of NH3:NO in 1 to 1.5 mole ratios
will reduce NO* emissions between 60 and 70%. Using a mole ratio of 0.5 will give
NO* reductions of approximately 40%.
Ammonia is mixed with a carrier gas of air or steam to provide good mixing of
NHj with the flue gas. This ammonia mixture is injected at various points in the
convection and superheater sections of the boiler (Figure 7-9). Multipoint injection
grids are used (to inject ammonia) to compensate for varying temperatures in the
convection and superheater sections. Temperature differences result from changing
electric loads to the generator.
Ammonia
injection
Figure 7-9. Exxon Thermal DeNO* process.
7-14
-------
NO, emissions in oil, gas, and small coal fired boilers can be reduced more than
60% by using ammonia injection. This process could be used to supplement the
NO, emission reductions accomplished by combustion modifications, when very low
levels of NO, are required. Strict NO, emission regulations adopted by the Califor-
nia Air Resources Board and the South Coast Air Quality Management District of
California require very low NO, emissions from utility boilers. The Los Angeles
Department of Water and Power is planning to install the Exxon Thermal DeNO,
process on their Haynes No. 4 unit, a 227 MW generator. The capital investment
to retrofit the Thermal DeNO, process is estimated to be approximately 18 to
25 $/kW. Operating and maintenance costs are estimated at 6 to 9 $/kW. Both
cost estimates are on a 1979 cost basis (EPA, Winter 1979-1980).
Selective Catalytic Reduction
The Selective Catalytic Reduction (SCR) is a dry process used to reduce NO, emis-
sions from fossil-fuel-fired boilers. The SCR process has been used extensively in
Japan to achieve a 90% reduction in NO, emissions (Mobley, 1979). This process is
based on the preferential reaction of NHS with NO, rather than with SO2 in the
flue gas. The reactions are expressed as:
(Eq. 7-2) 4NH3 + 4NO + O2 -4N2 + 6H2O
and
(Eq. 7-3) 4NHS + 2NO2 + O2 - 3N2 + 6H2O
Equation 7-2 represents the predominant reaction occurring since 95% of NO,
emissions in combustion flue gas are in the form of NO.
7-15
-------
This process involves injecting NH3 into the flue gas and passing this mixture
through a catalytic reactor (Figure 7-10). NO, emissions are reduced to harmless
molecular nitrogen (N2) and water vapor (H2O). Ammonia is injected on an
NHS:NO mole ratio of 1:1 attaining a 90% NO, emission reduction with less than
20 ppm NH3 leaving the reactor (Mobley, 1979).
Ammonia
injection
Catalytic
reactor
To air
preheater
Figure 7-10. Selective Catalytic Reduction (SCR) process.
The optimum temperature for successful NO, reduction in the catalytic reactor is
between 300 and 400 °C. The reactor is usually located between the boiler
economizer and air preheater. A bypass around the economizer is used when
temperatures begin to fall below 300 °C. The reactor can be located before or after
the baghouse or electrostatic precipitator used to collect paniculate matter.
A number of materials have been used for catalysts. Initially, a platinum metal
on an alumina (A12O3) support was used for NOX control on gas fired boilers.
Sulfur oxides, particularly SO3 and SO2, poison the alumina. Other catalysts must
be used which resist SO, deterioration when burning fuel containing sulfur.
Titanium dioxide (TiO2) and vanadium (V2OS) catalysts are resistant to SO, attack.
Therefore, most catalysts on oil and coal fired boilers contain TiO2 or V2O5
(Mobley, 1980). Other active metals including C, CO, Cr, Fe, Mn, Mo, Ni and W
have been used as catalysts. The oxides and sulfates of these metals have also been
7-16
-------
used as catalysts (Mobley, 1979). The exact compositions and constituents of most
catalysts are proprietary information.
Catalyst shape and reactor design vary, depending on the manufacturers' design
and the boiler application. Natural gas fired boilers use catalysts made of spherical
pellets, rings or cylinders that are packed in a. fixed bed reactor (Figure 7-11)
which is similar to the packed tower described in Chapter 4. Since high concentra-
tions of particulate matter will plug a fixed bed reactor, a moving bed reactor or a
parallel reactor is used on oil or coal fired boilers. In a moving bed reactor, the
catalyst is charged at the top of the reactor and moves down through the vessel
while the flue gas passes through the catalyst in a cross flow motion (Figure 7-12).
The spent catalyst is discharged at the bottom of the reactor. Particulate matter is
then removed by screening, and the catalyst is regenerated before being returned to
the reactor. Parallel flow reactors (Figure 7-13) were designed for use on coal fired
boiler flue gas containing a high particulate matter concentration. The flue gas
flow is parallel to the catalyst surface. Parallel catalyst shapes are either tubular,
honeycomb, or plates (Figure 7-14). Catalysts are usually made into unit cells that
are 12 inches square. These unit cells are stacked into the reactor to form the grids
or layers of the catalyst (Figure 7-15).
Figure 7-11. Fixed bed reactor.
7-17
-------
Charged catalyst
i
Figure 7-12. Mobile bed reactor.
Spent catalyst
Figure 7-13. Parallel flow reactor.
7-18
-------
Hi
N;
N
S3S
sn
^-
Figure 7-14. Parallel flow catalysts.
Figure 7-15. Unit cell detail of catalyst grid.
7-19
-------
The biggest problem with the SCR process is the formation of solid ammonium
sulfates (NH4)2SO4 and liquid ammonium bisulfate (Nt^HSO*) (Mobley, 1980).
This problem occurs when using the SCR process on high sulfur oil and coal fired
boilers. Ammonia reacts with SOS to form corrosive compounds of (NH^SO* and
NH4HSO4 that coat the air preheaters. Increased soot blowing in the furnace and
water washing of the air preheaters helps remove these materials.
In Japan, the SCR process has been used to reduce NO, on over fifty boilers or
furnaces burning gas or oil and is being installed on four coal fired utility boilers.
Expected startup of these facilities is mid to late 1981. In the U.S., a pilot plant
(0.5 MW) SCR process sponsored by the EPA has been installed at the Georgia
Power Company. The Southern California Edison Company is installing the SCR
process on an oil fired 100 MW unit at their Huntington Beach Station. A NO,
reduction of 90% is expected in order to meet California Air Resources Board'and
South Coast Air Quality Management District regulations.
The cost of the SCR process is estimated on a basis of a new 500 MW boiler
firing eastern bituminous coal with a heating value of 10,500 Btu/lb, a sulfur con-
tent of 3.5% and an ash content of 16%. The capital cost for 90% NO, reduction
(in 1979) is 42 $/kW and the annualized cost is 2.2 mills/kWh (Mobley, 1980).
Shell UOP Process
Shell UOP is a dry process that simultaneously removes both NO* and SO, emis-
sions. This process can also be designed to remove either compound separately
(Pohlenz, 1979).
The process uses a copper oxide (CuO) catalyst supported on alumina. These
catalysts are located in two or more parallel passage reactors as shown in Figure
7-16. Flue gas containing both NO, and SO, is introduced into the reactor where
the SO, reacts with copper oxide to form copper sulfate (CuSO4). At the same time
ammonia is being injected which reacts with the NO,. The copper sulfate, and to a
lesser extent the copper oxide, act as catalysts for the NO,-NHS reaction.
7-20
-------
Ammonia
injections
Figure 7-16. Shell UOP process.
The following reactions occur in the reactor:
(Eq. 7-4) CuO + V£ O2 + SO2 -
<£q- 7-5) 4NO + 4NHS + O2 CuS°4.4N2 + 6H2O
Catalyst
SO2
reduction
7-21
-------
When the reactor catalyst is saturated with CuSO4) the flue gas is redirected to a
fresh reactor and the spent catalyst is regenerated (Figure 7-17). Hydrogen is used
to regenerate the catalyst by reducing the CuSO4 to copper and producing a con-
centrated SO2 gas stream. The SO2 gas is then used to produce sulfuric acid or
elemental sulfur for commercial sale. The copper in the reactor is oxidized to CuO
and the process is ready to be put on line again. The reactions that take place in
the reactor during catalyst regeneration are:
(Eq. 7-6)
(Eq. 7-7)
CuSO4 + 2H2 - Cu + SOj + 2H2O
— CuO
Catalyst
regeneration
The Shell UOP process can be operated as a NO, emission reduction process by
eliminating the regeneration cycle. The process can be operated as a SO* emission
reduction process by eliminating the ammonia injection.
Hydrogen gas
Catalyst saturated
with CUSO4
Concentrated SO?
Figure 7-17. Catalyst regeneration (Shell UOP).
7-22
-------
This system can be designed to use moving packed beds, fixed packed beds, or
fluid systems. Shell UOP Inc. has developed a type of fixed bed reactor in which
the flue gas flows through open channels called a parallel passage reactor (Figure
7-18). The catalyst (CuO) is placed between thin layers of wire gauze. Flue gas
flows along the catalyst surfaces and the NOX and SO* gas diffuses into the catalyst.
This design reduces reactor plugging by paniculate matter. The NOr/SOx emission
reduction is achieved in less than one-half second (Pohlenz, 1979).
Figure 7-18. Detail of catalyst (Shell UOP).
The reactors are operated at temperatures between 350 and 450 °C during the
NO* reduction cycle. The pressure drop through the parallel passage reactor is
approximately 75 to 150 mm H2O depending on the gas flow rate and NO* reduction
desired. The ammonia injection rate is one mole of NH3 injected for one mole of
NO* reduced.
7-23
-------
The Shell UOP process has been operated on large boilers in Japan for both SO2
and NO* emission reduction. Units have operated to remove at least 90% SO2 and
70% NO, emissions (Mobley, 1980). EPA has sponsored a pilot plant demonstra-
tion unit at the Big Bend Station of the Tampa Electric Company in North
Ruskin, Florida. This demonstration unit (0.5 MW) is installed on a coal burning
boiler and is designed for 90% SO2 and 90% NO, removal efficiencies. The cost
estimates for a 500 MW utility boiler burning 3.5% sulfur and 16% ash coal are
given on a 1979 basis. The capital cost is 134 $/kW for SO2 and NO, removal and
the annualized cost is 6.4 mills/kWh (Mobley, 1980).
Wet NOX and SOX Processes
Wet processes for the simultaneous reduction of NO, and SO, emissions have been
developed and installed on large oil fired boilers in Japan. These processes use
absorbers to reduce both SO2 and NO, emissions. These processes are not cost com-
petitive compared to other flue gas treatment processes used in the United States at
this time. The estimated capital cost is 200 $/kW (SO2 and NO,) and annualized
cost is 11.3 mills/kWh (Mobley, 1980).
Other NO* Reduction Techniques
Glass Manufacturing
A project to reduce NO, emissions during the manufacture of glass is being con-
ducted by ORD's Industrial Environmental Research Laboratory in Cincinnati. In
conventional glassmaking processes, the silicon and other materials for producing
glass are heated in a furnace. Over 50% of the energy used to heat the furnace is
lost in exhaust gases. Experiments are under way to alter the furnace design so that
the exhaust gas is used to preheat the raw materials for glassmaking. Preheating
has a number of advantages in the glassmaking industry. It reduces the amount of
fuel needed in the furnace, the amount of air drawn into the combustion process,
the temperature required in the furnace to melt the silicon material, and the
length of time that the material must remain in the furnace. Reductions in fuel
use, air intake, and combustion temperature have resulted in the decrease of both
fuel-bound NO, emissions, and thermal NO, emissions. The costs that the glass
industry could save by using less fuel are an additional incentive to make use of the
preheating process (EPA, February 1980).
7-24
-------
Nitric Acid Production
About 70% of the nitric acid (HNO3) produced in the U.S. is used to manufacture
fertilizer. Other uses include the production of industrial explosives, separating
gold and silver, pickling steel and brass, and photoengraving. Nitric acid is pro-
duced by oxidizing ammonia (NH3). The oxidation is never totally complete,
however, and uncontrolled emissions from nitric acid plants are typically on the
order of 1000 to 3000 ppm NOX.
Several NO, control techniques are available and are being used. The IERL
Branch Laboratory in Edison, New Jersey has been developing a technique called
molecular sieve adsorption. NO, is removed by converting NO to NO2 and
adsorbing the NO2. This process results in NO, concentrations of less than 50 ppm
in the emission stream, and the NO2 which is collected can be used to produce
more nitric acid (EPA, Feburary 1980).
References
California Air Resources Board (GARB). 1978. Public Hearing on Petitions to
Review Rule 475.1 of the South Coast Air Quality Management District
May 25, 1978.
Chapman, K. S. 1979. NO, Reduction on Process Heaters with a Low NO, Burner.
Air Poll. Control Assoc. Meeting Paper, Cincinnati: 79-33.3.
Cotter, J. E. and Koppang, R. R. 1978. The Kurabo Process for Dry, Selective
Catalytic Reduction of NO,. Air Poll. Control Assoc. Meeting Paper, Houston-
78-28.5.
Environmental Protection Agency (EPA). Winter 1979-1980. NO, Control Review
EPA-IERL Vol. 5 No. 1.
Environmental Protection Agency (EPA). February 1980. Controlling Nitrogen
Oxides. EPA 600/8-80-004.
Environmental Protection Agency (EPA). January 1978. Control Techniques for
Nitrogen Oxides Emissions From Stationary Sources—Second Edition EPA
450/1-78-001.
Environmental Protection Agency (EPA). September 1974. Systems Analysis
Requirements for Nitrogen Oxide Control of Stationary Sources. EPA
650/2-74-091.
Environmental Protection Agency (EPA). July 1977. Proceedings of the Second
Stationary Source Combustion Symposium, Vol. II, Utility and Large Industrial
Boilers. EPA 600/7-77-073b.
Environmental Protection Agency (EPA). October 1977. Preliminary Environmen-
tal Assessment of Combustion Modification Techniques: Vol. I, Summary EPA
600/7-77-119a.
7-25
-------
Environmental Protection Agency (EPA). July 1977. Proceedings of the Second
V°L IV' ^^—tal Combustion
Environmental Protection Agency (EPA). June 1974. Field Testing: Application
of Combustion Modifications to Control NO, Emissions From Utility Boilers
EPA 650/2-74-066.
Environmental Protection Agency (EPA). March 1978. Environmental Assessment
of Stationary Source NO, Control Technologies: First Annual Report EPA
600/7-78-046.
Environmental Protection Agency (EPA). September 1975. NO, Combustion
Control Methods and Costs for Stationary Sources. EPA 600/2-75-046.
Electric Power Research Institute (EPRI). 1979. Proceedings: Second Annual NO,
Control Technology Seminar. EPRI/FP 1109-SR.
Jones G. D. 1980. Project Summary Selective Catalytic Reduction and NO, Con-
trol in Japan- A Status Report, October 1980. Submitted to EPA-IERL under
EPA Contract No. 68-02-3171, Task 10.
Lyon, R. K. and Tenner, A. R. 1978. Reducing NO, Emissions by Ammonia
Injection. Air Poll. Control Assoc. Meeting, Houston: 78-8.1.
Maxwell, J. D. and Burnett. 1980. Technical and Economic Evaluations of NO
Control Technology. Air Poll. Control Assoc. Meeting Paper, Montreal: 80-60.1.
Mobley, J. D. 1980. Assessment of NO, Flue Gas Treatment Technology
Symposium on Stationary Combustion NO, Control. Denver, Colorado, October
b-9, 1980.
Mobley, J. D. 1979. Flue Gas Treatment Technology for NO, Control U S
Environmental Protection Agency's Third Stationary Source Combustion Sym-
posium. San Francisco, California, March 5-8, 1979.
National Air Pollution Control Administration (NAPCA). 1969. Systems Study
of Nitrogen Oxide Control Methods for Stationary Sources. PB 192789.
National Air Pollution Control Administration (NAPCA). 1970. Control
Techniques for Nitrogen Oxide Emissions From Stationary Sources. AP-67.
Pohlenz, J. B. 1979. NO, and SO, Emissions Control with the Shell Flue Gas
Treating Process. Air Poll. Control Assoc. Meeting Paper, Cincinnati: 79-33.5.
Robinson, J. M. 1979. NO, Control by the Thermal DeNO, Process in a
Tertiary Oil Recovery Steam Generator. Air Poll. Control Assoc. Meeting Paper
Cincinnati: 79-33.4. v '
7-26
-------
Chapter 8
Control of Sulfur Oxide Emissions
Introduction
More than 25 metric tons of sulfur oxides (SO*) are emitted into the atmosphere
each year from manmade sources. These emissions can cause health problems,
reduce visibility, and contribute to the global acid rain problem (EPA, August 1980).
Sources of Sulfur Oxides
More than two-thirds of all manmade sulfur oxide emissions result from fossil fuel
combustion in electric generating plants (Figure 8-1). Approximately 8% of the
total sulfur oxides emitted are from industrial boilers. The largest noncombustion
sources are copper smelters, followed by petroleum refining operations. Other small
SO* sources include residential, commercial, and institutional heating furnaces, and
transportation vehicles (EPA, August 1980).
Industrial
boilers (8%)
Copper smelters (8%)
Petroleum refining (5%)
Transporation and others
(5%)
Residential, commercial,
institutional (5%)
Electric
generating
stations
(69%)
Source: EPA, August 1980.
Figure 8-1. Sulfur oxide emissions sources.
The majority of sulfur dioxide (SO2) emissions from utility boilers occurs when
coal is burned. When burned, most of the sulfur (S) contained in coal is oxidized
to SO2. Other fuels used by utilities such as natural gas and oil contain varying
amounts of sulfur. Much less SOZ is emitted from boilers burning natural gas than
from boilers burning high sulfur fuel oil. The consumption of coal for electricity
8-1
-------
generation is expected to increase annually from 475 million metric tons in 1975,
to between 600 and 1000 million metric tons in the year 2000 (EPA, August 1980).
The SO2 emissions from these sources are expected to increase from 20 to 41
million metric tons per year. Therefore, SOr emissions must be controlled if the
current SO2 ambient air quality levels are expected to be maintained or improved.
Approximately 95% of all sulfur oxides emitted from combustion processes are
in the form of sulfur dioxide (SO2). SO2 is a colorless gas that when cooled and
liquified is used as a bleach, disinfectant, refrigerant, or preservative. The
remaining sulfur oxide emissions are in the form of sulfur trioxide (SO,), sulfates
(SO4=), sulfites (SOsa) or sulfuric acid (H2SO4). SO2 can form sulfates when oxidized
in the atmosphere. Sulfur trioxide is not a stable compound and in the presence of
water forms sulfuric acid, a component of acid rain (EPA, August 1980).
Health Effects
High concentrations of sulfur oxides cause breathing difficulty. SO2 is very soluble
and is absorbed in the nose and upper air passages during breathing. This can
cause a choking effect known as pulmonary flow resistance. The degree of
breathing difficulty is directly related to the amount of SO2 in the ambient air.
Chronic coughing and mucus secretion may result from repeated exposure to high
levels of SO2. The young, elderly, and individuals with chronic lung or heart
diseases are most susceptible to the adverse effects of sulfur oxides (EPA August
1980).
Visibility
Atmospheric visibility is reduced by the presence of small suspended particles in
ambient air. Estimates of sulfate concentration in the total suspended particulate
matter range from 30 to 50% (EPA, August 1980). Analysis of ambient air quality
data in the South Coast Air Basin give estimates of sulfate and nitrate levels as
high as 30% of the TSP concentration during air pollution episodes in southern
California (Beachler, 1978). Various studies have shown a reduced visibility
occurring in the northeastern United States and other areas in the past 25 years.
EPA estimated that in some areas on a hot summer day, visibility has decreased
from 15 to 8 miles (EPA, August 1980).
Acid Rain
Acid rain primarily consists of sulfuric acid (H2SO4) and nitric acid (HNOS).
Sulfuric acid (H2SO4) comprises 40 to 60% of acid rain, depending on regional
emission patterns (EPA, August 1980). Rain in the northeastern United States
averages between 10 and 100 times the acidity of normal rainwater. More than 90
lakes in the Adirondack Mountains in New York no longer contain fish due to the
increased acidity of the lake water. In some cases, this acidity has caused toxic
metals in the lake beds and surrounding soils to be released into the water, conse-
quently killing fish (EPA, August 1980). Studies indicate damage to agricultural
crops, statues and monuments has occurred due to acid rain.
8-2
-------
New Source Performance Standards
EPA has promulgated New Source Performance Standards (NSPS) for fossil-fuel-
fired steam generators (FFFSG) with heat input greater than 250 x 108 Btu/hr (73
MW thermal). These standards establish sulfur oxide emission limits for utility
boilers. There are two standards: one for steam generators for which construction
commenced after August 17, 1971; the other for steam generators for which con-
struction commenced after September 18, 1978 (Table 8-1). The term "construc-
tion commenced" is defined in the Clean Air Act and does not necessarily mean
that actual physical construction was initiated. It may be interpreted as the date
that a permit process is initiated at the air pollution control agency. These stan-
dards also specified the degree of SO2 reduction required for utility boilers.
Table 8-1. New Source Performance Standards for SO, emissions from
fossil-fuel-fired steam generators rated greater than 250 x 10*
Btu/hr or 73 MW (thermal) heat input.
SO, emissions
Subpart D; new
sources after
August 17, 1981
Subpart Da; new
sources built after
September 18, 1978
Metric units
(ng/J)
340
520
340
and 90% SO, reduc
unless emissions are
86
520
and 90% SOt reduc
unless SOt emissions
260
then 70% SO2 redui
English units
db/10' Btu)
0.8
1.2
0.8
tion is required
less than
0.2
1.2
tion required
are less than
0.6
:tion is required
Fuel
Liquid
Solid (coal)
Liquid or gaseous
Solid (coal)
The NSPS for steam generators built after September 18, 1978 require flue gas
scrubbing to reduce SO2 emissions. In generators burning coal, the SO2 emissions
must be less than 1.2 lb/106 Btu. Also a 90% reduction of SO2 emissions is
required. If the emissions are less than 0.6 lb/106 Btu (coal fired units) then only
70% SO2 emission reduction is required. The purpose of this "two part" standard is
to encourage the use of new technology for reducing SO2 emissions from boilers
using low sulfur coal (JAPCA, 1979). Generators burning low sulfur coal generally
have SO2 emissions less than 0.6 lb/106 Btu. For oil fired utility boilers the SO2
emissions must be less than 0.8 lb/106 Btu and 90% SO2 emission reduction
(scrubbing) is required; unless the emissions are less than 0.2 lb/108 Btu. In this
case, no scrubbing is required.
For example, a power plant (>250x 106 Btu/hr) cannot have SO2 emissions
exceeding 1.2 lb/106 Btu. In addition, the plant is required to reduce the SO2
emissions by 90%. If the plant's emissions are less than 0.6 lb/106 Btu then only
70% scrubbing is required. It is possible that the SO2 emissions will fall between
0.6 and 1.2 lb/106 Btu. In this case, 90% scrubbing is required. It is also possible
8-3
-------
to meet the NSPS standard by scrubbing 85% if and only if the plant's emissions
do not exceed 0.6 lb/106 Btu. A useful graphical representation of the 1978 NSPS
for sulfur dioxide emission limitations from fossil-fuel-fired steam generators can be
found in JAPCA (1980).
SOg Emission Reduction Technology
Sulfur oxide emissions from fossil-fuel-fired combustion sources can be reduced by
five techniques: fuel desulfurization, combustion of coal and limestone mixtures,
coal gasification, coal liquefaction, and flue gas desulfurization.
The EPA's Office of Research and Development is providing research and
development projects on these technologies through grants, contracts, and
cooperative agreements.
Fuel Desulfurization
One of the most straightforward ways to reduce SO2 emissions from combustion
sources is by burning fuel containing less sulfur. This might involve using low
sulfur coal, low sulfur fuel oil, or natural gas instead of a high sulfur coal. The use
of low sulfur coal (usually supplied from western U.S.) can reduce SO2 emissions.
In some cases, low sulfur coal has been used to meet State and Federal (1.2 lb/106
Btu) air pollution regulations. Low sulfur coal generally contains less than 1%
sulfur; high sulfur coal contains between 3 and 5% sulfur. Low sulfur coal is more
expensive and less available (in the eastern U.S.) than high sulfur coal. However,
the sulfur content can be reduced by a process called fuel desulfurization.
Fuel desulfurization removes or reduces the sulfur content of coal before it is
burned. Since coal is projected to be the major domestic fuel to meet future energy
demands, fuel desulfurization can be a very useful method for reducing SO2 emis-
sions. Coal contains sulfur in two forms: mineral sulfur in the form of inorganic
pyrite and organic sulfur which is chemically bound to the coal. The amount of
inorganic (and organic) sulfur in coal varies from 10 to 90%. Mineral sulfur can be
removed by physical coal cleaning, but organic sulfur requires chemical cleaning.
Physical coal cleaning depends on the differences in density of both coal and
the impurities. Coal is crushed, washed, and then separated by settling processes
using cyclones, air classifiers or magnetic separators. Approximately 40 to 90% of
the pyritic sulfur content can be removed by physical coal cleaning (EPA, August
1980). Its effectiveness depends on the size of pyritic sulfur particles and the
amount of pyritic sulfur contained in the coal.
Chemical coal cleaning methods that reduce the organic bound sulfur are cur-
rently under development. The Industrial Environmental Research Laboratory
(IERL) of the EPA is involved in two chemical coal cleaning technologies:
microwave desulfurization and hydrothermal desulfurization.
In microwave desulfurization the coal is crushed, then heated for 30 to 60
seconds by exposure to microwaves. Mineral sulfur selectively absorbs this radiation
forming hydrogen sulfide gas (H2S). The H2S is usually reduced to elemental sulfur
8-4
-------
by the Claus process. Another microwave process adds calcium hydroxide
[Ca(OH)2] to crushed coal. The organic sulfur converts to calcium sulfite (CaSOs)
when exposed to this radiation. The coal is washed with water to remove the
CaSOj and other impurities. As much as 70% of the sulfur can be removed by the
microwave process (EPA, August 1980).
Hydrothermal desulfurization, developed by Battelle Laboratories in Columbus,
Ohio, is being refined through experiments sponsored by EPA's Office of Research
and Development (ORD). Coal is crushed and mixed with a solution of sodium and
calcium hydroxides [NaOH and Ca(OH)2]. When this mixture is heated to 275 °C
in a pressurized vessel, most of the pyritic sulfur and 20 to 50% of the organic
sulfur is converted to sodium and calcium sulfites (Na2SO, and CaSO3XEPA,
August 1980). The coal is rinsed to remove the sulfites and the water is processed
to recycle the sodium and calcium hydroxides. This process is an expensive but
effective method for removing sulfur from coal. ORD is directing current efforts
toward reducing high process costs and finding alternative methods for drying the
coal and recovering the sodium and calcium hydroxides.
Combustion of Coal and Limestone Mixtures
Sulfur oxides can be removed by burning coal and limestone mixtures in a boiler.
Two promising burning technologies currently under development are: Jluidized
bed combustion and the use of limestone coal pellets as fuel.
In the fluidized bed combustion process, a grid supports a bed of coal and
limestone (or dolomite) in the firebox of the boiler. Combustion air is forced
upward through the grid suspending the coal and limestone bed in a fluid-like
motion. Natural gas is used to ignite the pulverized coal. Once the coal is ignited,
the gas is turned off. The sulfur in the coal is oxidized to SO2 and consequently
combined with the limestone to form calcium sulfate (CaSO4). The CaSO4 and
flyash particulate matter are usually collected in a baghouse or electrostatic
precipitator (see APTI Course 413 Student Manual). A 100,000 Ib steam/hr
fluidized bed boiler is now operating at Georgetown University, Washington, DC.
The sulfur dioxide emissions from this unit have been reduced 75 to 90% (Milliken
and Young, 1981). The nitrogen oxide emissions were reported to be less than
0.7 lb/106 Btu due to low combustion temperatures (815 to 870°C). Other demonstra-
tions of fluidized bed combustion are being conducted at several sites including use
of a 30 MW boiler at Rivesville, West Virginia (EPA, May 1979). This technology
is expected to be available for commercial industrial boilers in the early 1980's.
Another way to reduce SO2 emissions from combustion processes is by burning
pellets made of limestone and coal. ORD research has shown that SO2 emissions
from conventional stoker and pulverized coal fired boilers can be reduced by
burning coal and limestone pellets. Pellets are made by pulverizing coal and
limestone and adding a binder forming small, cylinder shaped pellets which are
about half the size of a charcoal briquette. These consist of approximately two-
thirds coal, and one-third limestone. As the pellet burns, the calcium in the
limestone absorbs the SO2 generated by coal combustion and forms calcium sulfate.
8-5
-------
Calcium sulfate emissions are collected in a baghouse or electrostatic precipitator.
SO2 emissions have been reduced on a pilot scale boiler by as much as 80%
(Mobley and Linn, 1981).
Coal Gasification
Over 70 different processes have been developed for producing a combustible gas
from coal. Three basic steps are common to all coal gasification processes:
pretreatment, gasification, and gas cleaning. Coal pretreatment involves coal
pulverizing and washing. The pulverized coal is gasified in a reactor with limited
oxygen. Gasification produces either a low, medium, or high Btu gas by applying
heat and pressure or by using a catalyst to break down the components of coal
(EPA, August 1980). The gas produced contains carbon monoxide (CO), hydrogen
(H2), carbon dioxide (CO2), water (H2O), methane (CH*), and contaminants such
as hydrogen sulfide and char. Low and medium Btu gas contains more CO and H2
than high Btu gas which contains a higher CH* content. Methane gas produces
more heat when burned (see Table 3-1). The sulfur in the coal is converted to H2S
during gasification. H2S is removed during the gas cleaning step, generally by a
scrubbing process. H2S is converted to elemental sulfur by partial oxidation and
catalytic conversion (EPA, August 1980). The synthetic gas produced is sulfur free
and can be burned without releasing harmful pollutants. Commercial operation of
coal gasification units is expected in the mid 1980's.
Goal Liquefaction
A process for changing coal into synthetic oil is called coal liquefaction. Coal
liquefaction is similar to coal gasification. Two basic approaches for liquefaction
are used. One involves using a gasifier to convert coal to carbon monoxide,
hydrogen and methane; followed by condensation to convert the gases to oils. The
second approach uses a solvent or slurry to liquify pulverized coal and then
processes this liquid into a heavy fuel oil. Some processes produce both a synthetic
gas and synthetic oil.
Hydrogen is used to convert sulfur in the coal to hydrogen sulfide gas. Hydrogen
sulfide is partially oxidized to form elemental sulfur and water. More than 85% of
the sulfur is removed from coal by liquefaction (EPA, August 1980). EPA is cur-
rently working on improving these processes. Commercial operation of liquefaction
processes is expected in the early 1990's.
Flue Gas Desulfurization
Flue gas desulfurization (FGD) is the most popular technology used for controlling
sulfur oxide emissions from combustion sources. FGD technology is also used to
reduce SOr emissions from copper smelters. In this method SO* gaseous emissions
are generally removed by a post-combustion absorption process. This process uses
an absorbent in a vessel and produces a sludge that must be discarded.
8-6
-------
FGD scrubbing processes can either be wet or dry. Wet scrubbing processes use 3
liquid absorbent to absorb the SO2 gases. Wet scrubbing can be further categorized
into nonregenerable and regenerable processes. Nonregenerable processes produce
a sludge that must be disposed of properly; sometimes referred to as throw away
FGD processes. Regenerable processes generate a usable product in addition to
removing sulfur oxides. Regenerated products include elemental sulfur, sulfuric
acid, or gypsum wall-board. Dry scrubbing processes use a dry spray to absorb SO2
gas and form dry particles. These dry particles are collected in a baghouse or elec-
trostatic precipitator.
More than 50 different FGD processes have been developed. The majority of
these systems operating in Japan and the United States are capable of reducing the
sulfur oxide emissions from utility boilers by at least 90%. This chapter will
address the most important processes used in the U.S. today.
Nonregenerable FGD Processes
Nonregenerable FGD processes are those which generate a sludge or waste product
as a result of SO2 emission reduction. The sludge must be disposed of properly in a
pond or landfill. The three most common nonregenerable processes used in the
U.S. today are lime scrubbing, limestone scrubbing and double alkali scrubbing.
Approximately 75% of the current FGD installations in the U.S. are either lime or
limestone scrubbing (EPA, March 1978). However, this number could change with
additional use of regenerable processes to eliminate sludge disposal problems.
Lime Scrubbing
Lime scrubbing uses an alkaline slurry made by adding lime (CaO), usually 90%
pure, to water. The alkaline slurry is sprayed in an absorber and reacts with the
SO2 in the flue gas. Calcium sulfite (CaSOs) and calcium sulfate (CaSO4) salts are
formed in the reaction and are removed as sludge. The sludge produced can be
stabilized to produce an inert landfill material or can be stored in sludge ponds.
A number of chemical reactions take place in the absorber. SO2 can be absorbed
in the water and form sulfite (SOf) and sulfate (SO^) ions. Lime (CaO) is slaked
with water to produce a calcium slurry of CaO and H2O or calcium hydroxide
[Ca(OH),]. The calcium hydroxide/water mixture is a solution containing calcium
ions (Ca++) and hydroxide ions (OH"). Calcium ions combine with sulfate and
sulfite ions to produce a calcium sulfite and calcium sulfate sludge. The basic reac-
tions occurring are:
S02+CaOH + ^ H20 - CaSO3 + H2O
S02 + CaOH + V$ 02 + V$ H20 - CaSO4 + H2O
Lime, in a stoichiometric ratio of 1.0 to 1.1, will remove 1 mole of SO2 gas.
8-7
-------
System Description
The equipment necessary for SO2 emission reduction comes under four operations:
1. Scrubbing or absorption: accomplished with scrubbers, holding tanks, liquid
spray nozzles, and circulation pumps.
2. Lime handling and slurry preparation: accomplished with lime unloading and
storage equipment, lime processing and slurry preparation equipment.
3. Sludge processing: accomplished with sludge clarifiers for dewatering, sludge
pumps and handling equipment, sludge solidifying equipment.
4. Flue gas handling: accomplished with inlet and outlet ductwork, dampers,
fans, and stack gas reheaters.
The individual FGD systems vary depending on the FGD vendor and individual
plant installations. For example, some FGD systems have used a scrubber and
absorber for both paniculate and SO2 emission control (Figure 8-2). This system is
used at Pennsylvania Power's Bruce Mansfield 1 and 2 in Shippingport, Penn-
sylvania. This FGD system is designed to remove 99.8% paniculate matter and
92.1% SO2 emissions on an 825 MW coal fired utility boiler (1650 MW for both
units). Bruce Mansfield 1 was installed in late 1975. The system consists of six
parallel scrubber and absorber trains. Each train (Figure 8-2) consists of a variable-
throat venturi scrubber, a wet induced draft fan, and a fixed-throat venturi scrub-
ber (EPA, January 1979).
Flue gas from boiler
Paniculate
emission
removal
SO2 emission
removal (majority)
Variable-throat
venturi scrubber
Fixed-throat
venturi scrubber
Figure 8-2. Scrubber and absorber train for lime scrubbing system.
8-8
-------
Other FGD systems have been installed that consist of an electrostatic
precipitator for particulate emission control and a mobile bed absorber for SO2
emission control (Figure 8-3). This system was installed at Louisville Gas and Elec-
tric Company's Cane Run 190 MW power station, Louisville, Kentucky. The system
is designed for 99.0% particulate emission control and 85 to 89% SOj emission
control. The startup of this system was in August, 1976 (EPA, January 1979).
Particulate
emission
removal
Electrostatic precipitator
Mobile bed scrubber
Figure 8-3. Electrostatic precipitator and mobile bed scrubber train
for lime scrubbing system.
8-9
-------
A listing of selected lime scrubbing systems installed in the U.S. is given in Table
8-2. This listing contains the unit rating for the power plant, the FGD system
emission control efficiency, startup dates and sludge disposal method. Additional
installations can be found in Electric Utility Generating Units-Flue Gas
Desulfurization Capabilities as of October 1978 (EPA, January 1979).
Table 8-2. Selected lime scrubbing FGD systems.
Station
Kentucky Utilities;
Green River
1, 2 and 3
Pennsylvania Power
Co.; Bruce
Mansfield 1 and 2
Columbus and
Southern Ohio
Electric Co. ;
Conesville 5 and 6
Louisville Gas
and Electric;
Cane Run 4
Louisville Gas
and Electric;
Cane Run 5
^™— •
Size
(MW)
64
825
825
411
375
178
183
Initial
startup
date
8/75
12/75 (unit 1)
10/77 (unit 2)
1/77
6/78
8/76
12/77
New or
retrofit
••— ^— — ••— •••
Retrofit
New
New
Retrofit
Retrofit
Paniculate
emission
control
—
Variable-throat
venturi
scrubber
Electrostatic
precipitator
Electrostatic
precipitator
Electrostatic
precipitator
SO,
emission
control
—
Fixed-throat
venturi
scrubber
Turbulent
contact
absorber
(mobile bed)
Mobile bed
absorbers
Spray towers
Coal
sulfur
content
(%)
3.7
2.4
4.2 to 5.1
3.8
3.8
SO, emission
removal
efficiency
(%)
80
92-95
89.6
86-89
85
Sludge
disposal
-
Stabilized
sludge in an
off-site
dammed
resevoir
Stabilized sludge
disposed in an
outside landfill
Stabilized sludge
in an on-site pond
Stabilized sludge
n an on-site pond
Sludge produced from a lime scrubbing FGD system must be discarded properly
It can be stored m a settling pond or dewatered to take up less space in a landfill
Clanfiers, centrifuges and vacuum filters accomplish dewatering. Occasionally
chemicals are used to make the sludge into a cement-like material. The Bruce
Mansfield plant uses a material called Calcilox (developed by Dravo Corporation)
The cement-like sludge is permanently stored in an earthen dammed reservoir
7 miles from the plant site (EPA, January 1979).
Operating Experience
Early lime FGD systems were plagued with a number of operational and
maintenance problems. One problem was that spray nozzles and mist eliminators
became clogged with hardened scrubbing slurry. Scale buildup (CaSO4) on the
spray nozzles limited scrubber operation time. New spray nozzle design and careful
control of the recirculated slurry have reduced internal scrubber scaling (EPA
Technology Transfer, Second Progress Report for Lime/Limestone Scrubbing)
Mist eliminator clogging was a major problem at the Bruce Mansfield plant.
Scaling on the mist eliminators caused excessive pressure drops. A wash system was
devised to reduce excessive scale buildup. Other problems with occasional mist
8-10
-------
carry-over caused scale buildup on flue linings and caused acid rainout in areas
near the stack. These problems have been eliminated by the careful design of mist
eliminators and by monitoring pressure drop across them (EPA, January 1979).
Another problem concerned stack gas reheaters. Stack gas is reheated to avoid
condensation and corrosion to ductwork and the stack; and to enhance plume rise
and pollutant dispersion. Reheat is accomplished by using steam coils in the stack
or by using hot air supplied by auxiliary oil heaters in the stack. Some reheater
failures were caused by acid attack to reheater components. Other reheaters
vibrated too much causing structural deterioration. These problems have been
eliminated by changes in reheater design (EPA, January 1979).
Corrosion of scrubber internals, fans and ductwork, and stack linings have been
reduced by using special materials such as rubber or plastic coated steel and by
carefully controlling slurry pH with monitors. Additional operation and
maintenance problems and solutions are found in Proceedings: Symposium on
Flue Gas Desulfurization, Volume I and II (EPA, March 1978, July 1979, and
1981).
Limestone Scrubbing
Limestone scrubbing uses an alkaline slurry from limestone (CaCOs) in an absorber
to react with SO2 in the flue gas. Calcium sulfite (CaSO3) and calcium sulfate
(CaSO4) salts are formed in the reaction and are removed as sludge. Two major
differences between lime and limestone scrubbing are their uses of feed preparation
equipment and their liquid-to-gas ratios. Even with these differences, these
processes are so similar that an FGD system can be set up to use either lime or
limestone to absorb SO2 gas.
Limestone scrubbing's process chemistry is also very similar to lime scrubbing's.
Limestone (CaCO3) is slaked with water to form aqueous CaCO3 and is sprayed in
the absorber. Sulfite and sulfate ions are produced as SO2 gas contacts the water.
These ions combine with calcium ions to produce calcium sulfite and calcium
sulfate sludge. The basic reactions are:
SO2 + CaCO3 + H2O + O2-CaSO3 + H2O + CO2 + O2
SO2 + CaCO3 + H2O + i/4 O2 — CaSO4 + H2O + CO2
System Description
The equipment necessary for SO2 absorption is the same as for lime scrubbing
except in the slurry preparation. The limestone feed (rock) is reduced in size by
crushing it in a ball mill. Limestone is sent to a size classifier. Pieces larger than
200 mesh are sent back to the ball mill for recrushing. Limestone is slaked with
water in a slurry supply tank. Limestone is generally a little cheaper than lime,
making it very popular for use in FGD systems.
8-11
-------
Limestone FGD systems vary, depending on the FGD vendor and individual
plant installations. The Kansas City Power and Light La Cygne power station uses
a venturi scrubber for particulate emission control and a sieve tray absorption
tower for SO2 emission control. The 820 MW plant uses eight scrubbing modules;
one venturi scrubber and one absorber per module as shown in Figure 8-4. The
Northern States Power Co. Sherburne Station No. 1 and 2 uses a scrubbing system
consisting of 12 scrubber and absorber modules. Other systems use electrostatic
precipitators for particulate emission control and absorbers for SOZ emission
control. A listing of selected limestone scrubbing U.S. installations is given in
Table 8-3.
Particulate
emission
removal
Venturi scrubber
Sieve tray tower
Figure 8-4. Limestone scrubbing system.
8-12
-------
Table 8-3. Selected limestone scrubbing FGD systems.
Station
Kansas City Power
and Light Co.;
La Cygne 1
Northern Scales
Power Co.;
Sherburne 1
and 2
Arizona Public
Service Co. ;
Choila 1
Arizona Public
Service Co. ;
Choila 2
Central Illinois
Light;
Duck Creek 1
Tennessee Valley
Authority;
Widows Creek
Texas Utilities;
Martin Lake
1 and 2
Kansas Power and
Light;
Lawrence 4
Kansas Power and
Light;
Lawrence 5
Size
(MW)
820
680
680
119
250
400
550
750
(each)
125
420
Initial
startup
date
2/73
3/76
1/77
10/73
6/78
8/78
5/77
10/77
1/76
4/78
New or
retrofit
New
New
Retrofit
New
New
Retrofit
New
Retrofit
New
Paniculate
emission
control
Venturi
scubber
Rod deck
venturi scrubber
Cyclone and
flooded disc
venturi scrubber
Flooded disc
scrubber
Electrostatic
precipitator
Variable-throat
venturi scrubber
Electrostatic
precipitator
Rod deck
venturi scrubber
Rod deck
venturi scrubber
SO,
emission
control
Sieve
tray absorber
Spray
tower
Spray
tower
Packed
tower
Rod deck
venturi scrubber
Multigrid
absorption tower
Packed
tower
Spray
tower
Spray
tower
Coal
sulfur
content
(%)
5.0
0.8
0.4-1.0
0.4-1.0
2.5-3.0
3.7
1.0
0.9
0.55
0.55
SO, emission
removal
efficiency
(%)
80
50-55
50-60
75
92
85-94
74
71
96-98
65
Sludge
disposal
Settling pond
Settling pond
Unstabilized sludge
in an on-site pond
Unstabilized sludge
in an on-site pond
Unstabilized sludge
in an on-site pond
Unstabilized sludge
in an on-site pond
Stabilized dewatered
sludge in an on-site
landfill
Unstabilized sludge
in an on-site pond
Unstabilized sludge
in an on-site pond
Operating Experience
Early limestone FGD systems had scrubber operating problems similar to lime
scrubbing systems. Plugged and clogged nozzles, scrubber internals, and mist
eliminators resulted from inefficient SO2 absorption by limestone in the scrubber.
Increased absorption efficiency is achievable at high pH values since more alkali is
available to dissolve the SO2 gas. However, scaling problems will occur if the scrub-
ber is operated at very high pH values (>9.0). The pH levels must be maintained
by carefully controlling lime or limestone feed and mixing with water to prevent
inefficient system operation. If the pH gets too low (<5.0) the removal efficiencies
are low; and at high pH levels (>9.0), scaling of scrubber components will occur.
As can be seen from Tables 8-2 and 8-3 the SO2 removal efficiencies for various
lime and limestone FGD installations range from 50 to 92 %. These FGD systems
were designed to meet existing air pollution regulations. Lime and limestone FGD
8-13
-------
systems are capable of removing SO2 with efficiencies in excess of 90% (EPA
March 1978; EPA 600/7-78-032a). The addition of small amounts of soluble'
magnesium (< 1% by weight) to the scrubber liquor can greatly increase SO,
removal efficiencies to as high as 99% (EPA, March 1978; EPA 600/7-78-032a).
Magnesium is added in the form of magnesium oxide, magnesium sulfate or
dolomitic lime (used in lime scrubbing systems). Magnesium compounds are more
soluble than calcium compounds and react rapidly with SO2.
EPA is currently working on a program that uses an additive of adipic acid to
limestone FGD systems. Adipic acid can increase SO2 removal efficiencies from
85% to as high as 97% (EPA, August 1980). Adipic acid is a crystalline powder
derived from petroleum. EPA experiments have shown that when a limestone slurry
reacts with SO2 in the scrubber, the slurry becomes very acidic. This acidity limits
the amount of SO2 that can be absorbed. Adding adipic acid to the slurry slightly
increases the slurry's initial acidity, but prevents it from becoming highly acidic
during the absorption of SO2. The net result is an improvement in the scrubbing
efficiency.
EPA research has shown that adipic acid can reduce the total limestone con-
sumption by as much as 15%, thus reducing operating costs. Adipic acid is non-
toxic (it is used as a food additive) and does not degrade the calcium sulfite and
calcium sulfate sludge. Adipic acid is not currently being used in a commercial
FGD installation. Full scale tests at an electrical generating station are in the
planning stage (EPA, August 1980).
Another scrubber operating problem occurring in lime and limestone FGD
systems is that calcium sulfite is formed as part of the sludge. Calcium sulfite settles
and filters poorly and can be removed from the scrubber slurry only in a semiliquid
or paste-like form. This semiliquid waste must then be stored in lined ponds. The
EPA's IERL group at RTF, NC is developing a way to solve this problem through
a process called forced oxidation.
In forced oxidation, air is blown into the scrubber slurry tank that contains
primarily calcium sulfite and water. The air oxidizes the calcium sulfite to calcium
sulfate.
CaSO3 + H20 + i/£ 02 -CaS04 + H2O
Calcium sulfate formed by this reaction grows to a larger crystal size than calcium
sulfite. As a result, the calcium sulfate is easily filtered, forming a drier and more
stable material which can be disposed of in a landfill. This material (CaSO4) can
also be used for cement, gypsum wallboard, or as a fertilizer additive.
Forced oxidation can also help control scaling problems to scrubber internals.
This process helps control scale by removing calcium sulfite from the slurry in the
form of calcium sulfate which is more easily filtered. This will prevent calcium
sulfites and calcium sulfates from being recirculated in the absorber.
8-14
-------
Double Alkali Scrubbing
Double or dual alkali scrubbing is a nonregenerable FGD process that uses two
alkali sprays to remove SO2 from combustion exhaust gas. A sodium solution
absorbs SO2 followed by absorbent regeneration. The spent absorbent is reacted
with a lime or limestone alkali slurry. Calcium sulfites and sulfates are precipitated
and discarded as sludge. The regenerated sodium scrubbing solution is returned to
the absorber loop. The double alkali process has reduced plugging and scaling
problems in the absorber since sodium scrubbing compounds are very soluble. FGD
systems are capable of 95% SO2 reduction.
Particulate matter is removed prior to scrubbing in the absorber by an elec-
trostatic precipitator or a venturi scrubber. This is done to prevent fly ash erosion
to the absorber internals and to prevent any appreciable oxidation to the sodium
solution in the absorber due to catalytic elements in the fly ash (EPA, March
1978). A typical dual alkali system is shown in Figure 8-5. SO2 removal in the
absorber occurs by reacting SO2 gas with a circulating sodium alkaline solution.
Electrostatic precipitator
Sieve tray tower
Figure 8-5. Dual alkali scrubbing system.
The sodium solution is usually a mixture of sodium carbonate (soda ash, Na2CO3),
sodium sulfite (Na2SO3), and sodium hydroxide (caustic, NaOH). During the reac-
tion the sodium solution is consumed to form sodium bisulfite (NaHSO3), sodium
sulfate (Na2SO4), and sodium bicarbonate (NaHCO3) salts. SO2 absorption can be
represented by the following simplified reactions:
8-15
-------
2Na2CO3 + SO2 + H2O-Na2SO3 + 2NaHCO3
NaHCO3 + SO2 - NaHSOj + CO2 !
2NaOH + SO2 - Na2SOs + H2O
Na,SOs + SO2 + H2O -2NaHS03
2NaOH + SO3 -Na2SO4 + H2O
O2-Na2SO4
After reaction in the absorber, spent scrubbing liquor is bled to a reactor tank
for regeneration. Sodium bisulfite and sodium sulfate are inactive salts and do not
absorb any SO2. Actually, it is the hydroxide ion (OH), sulfite ion (SO3) and car-
bonate ion (COS) that absorb SO2 gas. Sodium bisulfite and sodium sulfate are
reacted with lime or limestone to produce a calcium sludge and a regenerated
sodium solution:
2NaHSO3 + Ca(OH)2 - Na2SO3 + CaSO3 • ft H2O i + 3/2H2O
(lime) (sludge)
Na2SO4 + Ca(OH)2 -2NaOH + CaSO4 i
(lime) (sludge)
From the reactor, the slurry is pumped either to a clarifier or thickener where
precipitated solids (sludge) are separated from the scrubbing liquor. These solids
are dewatered by a vacuum filter and occasionally stabilized with a chemical or a
lime and fly ash mixture. Unstabilized sludge is discarded in a settling pond;
stabilized sludge in a proper landfill. Some sodium sulfate is unreacted (lost) in the
regeneration step. Additional sodium is added to the regenerated solution in the
form of soda ash or caustic soda. This regenerated absorbent is now ready to be
used again.
System Description
Currently, over 12 dual alkali processes are installed in the U.S. as FGD systems
on utility and industrial boilers. Louisville Gas and Electric Co. installed this
process on a 300 MW boiler at the Cane Run power station. Paniculate matter is
removed by an electrostatic precipitator before entering the absorber. The absorber
is a tray tower that is designed to remove 90 to 95% of the SO2 emissions generated
from burning 3.5 to 4.0% sulfur coal. A dual alkali system using three absorbers
was installed on three industrial boilers (total 165 MW) at the Arco Polymers plant
in Monaca, Pennsylvania. This system removed 86 to 93% of the SO2 emissions
generated from burning coal with up to 3% sulfur content. This system was
brought on line on August 16, 1980 and after a shakedown period has been
operating in compliance with State regulations of 0.75 lb/106 Btu (Reiners et al.,
1981). A listing of some selected dual alkali installations is presented in Table 8-4.
8-16
-------
Table 8-4. Selected dual alkali FGD systems.
Station
or
plant
Louisville Gas and
Electric; Cane
Run 6,
Louisville, KY
Southern Indiana
Gas and Electric
Co.; West
Franklin, Ind.
Arco Polymers;
Monaca, PA
Caterpillar tractor;
Joliet, IL
General Motors;
Parma, OH
Central Illinois
Public Service;
Newton, IL
Size
(MW)
SOO
500
165
34
64
575
Initial
startup
date
2/79
4/79
8/80
1974
1974
1979
Paniculate
emission
control
Electrostatic
precipitator
Electrostatic
precipitator
Electrostatic
precipicator
Mechanical
collector
—
Spray tower
SO,
emission
control
Tray tower
Two stage
disc absorber
Four stage
disc and doughnut
baffle absorber
2 Zurn dustraxtor
scrubbers
Tray tower
Mobile bed
absorber
Coal
sulfur
content
(%)
S. 5-4.0
45
S.O
3.2
2.5
4.0
SO,
removal
efficiency
(%)
90-95
85
86-9S
90
90
95
Operating Experience
As can be seen from Table 8-5, dual alkali scrubbing systems have been operating
for a number of years. Typical problems with pH control, mist eliminator liquid
carry-over, raw materials handling, and regenerated scrubber liquor pumping
operations have occurred. These problems have been rectified and installed systems
have been operating reliably for a number of years. The major benefit of the dual
alkali process is that by using a sodium solution as the absorbent, many scaling
problems that have plagued lime and limestone scrubbers can be eliminated.
However, double alkali scrubbing is more expensive than lime and limestone
scrubbing.
Regenerable FGD Processes
Regenerable FGD processes are wet scrubbing processes that remove SO2 from the
flue gas and generate some usable product. Regenerated products include elemen-
tal sulfur, sulfuric acid, or in the case of lime and limestone scrubbing, gypsum
wallboard. Regenerable processes used on industrial and utility boilers include
Wellman-Lord, magnesium oxide, and citrate. Approximately 4% of all operating
and committed FGD systems in the U.S. use the Wellman-Lord process; 2% use
the magnesium oxide process; and less than 1% use the citrate process (JAPCA,
1981). The major advantage of a regenerable process as opposed to a
nonregenerable process is the avoidance of a sludge disposal problem.
8-17
-------
Wellman-Lord
The Wellman-Lord process is a regenerable FGD process used to reduce SO2 emis-
sions from utility and industrial boilers and produce a usable product. This process
is sometimes referred to as the Wellman-Lord/Allied Chemical process; Allied
Chemical referring to the regeneration step.
SO2 is absorbed by aqueous sodium sulfite solution forming a solution of sodium
bisulfite. The solution is sent to an evaporator-crystallizer where sodium sulfite and
concentrated SO2 gas are produced. Sodium sulfite is recycled to the absorber and
concentrated SO2 is converted to a saleable product; usually sulfuric acid or
elemental sulfur. Flue gas from the boiler is pretreated prior to entering the
absorber. Paniculate matter is removed by an electrostatic precipitator, fabric
filter, or a wet scrubber. A small venturi scrubber removes any additional par-
ticulate matter, chlorides and sulfur trioxide. Flue gas is then sent to the absorber
where SO2 gas is absorbed according to the following reaction:
S02 + Na2SOs + H2O-2NaHS03
Some oxidation occurs in the absorber forming sodium sulfate which is unreactive
with SO2 gas:
Na2SO3 + i/£O2~Na2SO4
The formation of sodium sulfate depletes the supply of sodium sulfite available for
scrubbing. This can be made up by adding sodium carbonate to the scrubbing
slurry to combine with sodium bisulfite according to the following chemical reaction:
Na2CO3 + 2NaHSOs - 2Na2SO3 + CO21 + H2O
System Description
The Wellman-Lord process was installed at the Northern Indiana Public Service
Company's Dean H. Mitchell plant on a 115 MW boiler. The process equipment
includes: an electrostatic precipitator for removing paniculate matter; a venturi
scrubber for cooling flue gas and for removing SO3 and chlorides; an SO2 absorber;
an evaporator-crystallizer for regenerating the absorbent; and the Allied Chemical
process for reducing concentrated SO2 gas into elemental sulfur or sulfuric acid
(Figure 8-6). The absorber is a tray tower with three stages. SO2 gas is scrubbed
with a sodium sulfite solution at each stage. A mist eliminator removes entrained
liquid droplets from gas exiting the absorber. There is a direct-fired natural gas
reheat system in the absorber stack to reheat cleaned gas for good dispersion of the
steam plume.
8-18
-------
Regenerated absorbent
Electrostatic
precipitator
Concentrated SOj
Figure 8-6. Wellman-Lord system.
The solution (sodium bisulfite), collected at the bottom of the absorber,
overflows into an absorber surge tank. This solution is pumped through a filter to
remove any collected paniculate matter. A small side-stream is sent to a purge
treatment tank where sodium sulfate is removed. The solution is then pumped to
the evaporator for regeneration of the sodium sulfite solution.
The evaporator is a forced-circulation vacuum evaporator. Solution is recir-
culated in the evaporator where low pressure steam evaporates water from the
sodium bisulfite solution. When sufficient water is removed, sodium sulfite crystals
form and precipitate. Concentrated SO2 gas (95% by volume) is removed by the
steam. The sodium sulfite crystals form a slurry that is withdrawn continuously and
sent to a dissolving tank, where condensate from the evaporator is used to dissolve
the sodium sulfite crystals into a solution. This solution is pumped back into the
top stage of the absorber (EPA, Technology Transfer Capsule Report; EPA
625/2-77-011).
The following reaction takes place in the evaporator:
2NaHSO3 -Na2SO3 + H2O + SO21
(concentrated)
The water vapor is removed from the evaporator's overhead SO2/H2O vapors by
water-cooled condensers. The SO2 is compressed by a liquid ring compressor and
sent to the Allied Chemical SO2 reduction plant.
8-19
-------
Operating Experience
The Wellman-Lord process has been installed on two 350 MW coal fired boilers at
the Public Service of New Mexico San Juan power plant. This system was supplied
by Davy Powergas and is similar to the system supplied on the Mitchell power plant
in Indiana. The San Juan plant uses a five stage tray tower instead of the three
stage tray tower of the Mitchell plant. The Wellman^Lord process has also been
used to control SO2 emissions at Glaus tail gas plants^ sulfuric acid plants, and on
industrial boilers. A listing of selected installations is| given in Table 8-5. SO2
removal efficiency ranges from 85 to 90%, with a high of 98% on units installed in
Japan (EPA, March 1978; EPA 600/7-78-032b).
Table 8-5. Wellman-Lord installations in the United States.
Company and location
Units on stream:
Olin Corporation
Paulsboro, NJ
Standard Oil of California
El Segundo, CA
Allied Chemical Corporation
Calumet. IL
Olin Corporation
Curtis Bay, MD
Standard Oil of California
Richmond, CA
Standard Oil of California
El Segundo, CA
Standard Oil of California
Richmond, CA
Northern Indiana Public
Service Company
Gary, IN
Public Service Company
of New Mexico
Waterflow, NM
Public Service Company
of New Mexico
Waterflow, NM
Units in design or construction-*
Getty Oil Company
Delaware City, DE
Getty Oil Company
Delaware City, DE
Getty Oil Company
Delaware City, DE
Public Service Company
of New Mexico
Waterflow, NM
Public Service Company
of New Mexico
Waterflow, NM
Feed gas origin
Sulfuric acid plant
Claus plant
Sulfuric acid plant
Sulfunc acid plant
Claus plant
Claus plant
Claus plant
115-MW coal fired with 80%
load factor and recovery
capacity
375-MW coal fired boiler
system, San Juan
Station No. 1
375-MW coal fired boiler
system, San Juan
Station No. 2
60- MW mixed fuel boiler
system, Delaware City No. 1
60-MW mixed fuel boiler
system, Delaware City No. 2
60-MW mixed fuel boiler
system, Delaware City No 3
550-MW coal fired boiler
system, San Juan
Station No. 3
550-MW coal fired boiler
system, San Juan
Station No 4
Gas volume treated
(m'/s)
20.1
13.4
13.4
34.8
13.4
13.4
13.4
105.0
840.0
235.0
1121 0
(scfm)
43,000
28,000
28,000
74,000
28,000
28,000
28,000
223,000
1,780,000
520,000
2,400,000
Initial
startup
date
1970
1972
1973
1973
1975
1975
1976
1977
1978
1980
1981
*As of February, 1979
Source: EPA, February 1979.
8-20
-------
Magnesium Oxide Process
Magnesium oxide scrubbing is a regenerable FGD process used to remove SO, from
combustion exhaust gas. Magnesium oxide (MgO) slurry absorbs SO, and forms
magnesium sulfite. Magnesium sulfite solids are separated by centrifugation and
dried to remove moisture. The mixture is calcined to regenerate magnesium oxide
and produce concentrated SO, gas for production of sulfuric acid or elemental
sulfur.
Paniculate matter is removed from boiler exhaust by a precipitator or wet scrub-
ber prior to entering the absorber. Magnesium oxide slurry is sprayed and absorbs
SO2 according to the following simplified reactions:
Mg(OH), + 5H,0 + SO,-MgSOs»6H,0
MgS03.6H,0 + S02-Mg(HS03), + 5H,O
Mg(HSOs), + MgO -2MgSO, + H,O
2MgSO, + O, + 7H20-~2MgS04«7H,0
The aqueous slurry used for scrubbing contains the hydrated crystals of MgO,
MgSOs and MgSO4. A continuous side stream of this recycled slurry is sent to a
centrifuge where partial dewatering produces a moist cake. The liquor removed
from the crystals is returned to the main slurry stream. The moist cake is dried at
350 to 450 °F in a direct contact or rotary bed dryer. The dried cake is then sent to
a calciner where coke is burned at very high temperatures (1250 to 1S40°F) to
regenerate magnesium oxide crystals according to the following reactions:
Cake dryer
MgSOs«6H,O- — MgSO, + 6H,Ot
neat
MgS04-7H20- — MgS04 + 7H2Ot
neat
MgO regeneration in calciner
MgSO,- — MgO + SO2t
neat (concentrated)
r — COt
heat
CO
+ MgSO4- — CO2 + MgO + SO, t
heat (concentrated)
8-21
-------
System Description
The magnesium oxide process has been installed on a few power plant boilers
burning coal or oil in the U.S. A list of some installations is given in Table 8-6. A
test project was installed at Boston Edison's Mystic Station, a 155 MW plant that
burns oil. Since oil was burned, a low concentration paniculate matter was in the
flue gas and did not need to be removed before the gas entered the absorber.
Magnesium oxide slurry absorbed SO2 in a (cocurrent) venturi scrubber. SO2
scrubbing, slurry dewatering, and magnesium sulfite drying were done at the
Mystic station. Magnesium oxide regeneration and sulfuric acid production was
done at the Essex Chemical plant in Rumford, Rhode Island. Regenerated
magnesium oxide was returned to the Mystic station. The average SO2 removal was
approximately 92% (EPA, March 1978; EPA 600/7-78-032b). This test project was
terminated in 1974. This EPA publication gives the major operating problems and
solutions of three magnesium oxide scrubbing installations in the United States.
At the Philadelphia Electric Eddystone Station, Unit 1A, the MgO process was
installed on a 120 MW coal fired utility boiler burning 2.5% sulfur coal. Most of
the paniculate matter is removed by an electrostatic precipitator; the rest by a ven-
turi scrubber. Flue gas then enters a venturi rod scrubber where more than 90% of
the SO2 in the flue gas is removed. The MgO is regenerated in the Essex Chemical
plant in New Jersey. Operating problems and solutions are given in EPA
600/7-78-058b (EPA, 1978).
Table 8-6. Magnesium oxide installations in the United States.
Station
Boston Edison;
Mystic station
Potomic Electric
and Power;
Dickerson No. 3
Philadelphia
Electric;
Eddystone No. 1A
Philadelphia
Electric.
Eddystone No. IB
Size
(MW)
155
95
120
240
Startup
date
9/72
9/73
9/75
6/80
Fuel
Oil
Coal
(2.0% sulfur)
Coal
(2.5% sulfur)
Coal
(2.5% sulfur)
Paniculate
emission
control
-
Electrostatic
precipitator
Electrostatic
precipitator and
venturi scrubber
ESP and
venturi
SO,
emission
control
Venturi
scrubber
Venturi
scrubber
Venturi
rod scrubber
Venturi
rod scrubber
SO, emuiion
removal
efficiency
92
82-93
90
90
Comments
Project
terminated 1974
Project
terminated 1975
Citrate Process
The citrate process, a regenerable FGD process, was developed by the U.S. Bureau
of Mines. The citrate process uses sodium citrate and citric acid as buffering agents
to attain a higher solubility of the SO2 in an aqueous absorbent solution. The
chemistry of this process is very complex. The absorption of SO2 is pH dependent,
increasing with higher pH. SO2 forms H2SO3 when absorbed by water, resulting in
decreasing pH values. This creates a more acidic solution that inhibits additional
8-22
-------
absorption of SO2 gas. By using a buffering agent to prevent a pH drop, a substan-
tially higher amount of SO2 can be absorbed. One commercial installation exists at
this time, on a 60 MW industrial boiler at St. Joe Minerals in Monaca, Penn-
sylvania. Particulate matter is removed from the flue gas by an electrostatic
precipitator. Chlorides and sulfuric acid mist are removed from the flue gas by a
small venturi scrubber before entering the packed tower. SO2 is absorbed by a solu-
tion of sodium citrate, citric acid and sodium thiosulfate to produce sodium sulfite.
SO2 removal efficiency is approximately 90% (Madenburg et al., 1979). The solu-
tion (containing absorbed SO2) is reacted with hydrogen sulfide gas (HtS) in a
closed vessel to precipitate elemental sulfur and regenerate citrate solution. The
elemental sulfur precipitate is concentrated by air flotation into a sulfur slurry
which is separated from the regenerated solution. Sulfur slurry is heated to form
liquid sulfur and the solution is decanted. Hydrogen sulfide gas (used in regenera-
tion) is obtained either as a by-product of petroleum refining or produced on site
by reacting recovered sulfur with natural gas and steam.
The citrate process has some drawbacks. It is expensive compared to lime and
limestone scrubbing. Production of H2S from methane and sulfur can be a problem
due to the shortage of natural gas. Since the size of the absorption, regeneration,
and flotation equipment is large, boiler retrofitting is sometimes difficult.
Dry SO2 Control Systems
One promising new technology for reducing SO2 emissions from combustion sources
is using dry flue gas desulfurization (FGD). In dry FGD, the flue gas containing
SO2 is contacted with an alkaline material to produce a dry waste product for
disposal. This technology includes:
• injection of an alkaline slurry in a spray dryer with collection of dry particles
in a baghouse or electrostatic precipitator (ESP);
• dry injection of alkaline material into the flue gas stream with collection of
dry particles in a baghouse or ESP;
• addition of alkaline material to the fuel prior to combustion.
These technologies are capable of SO2 emission reduction ranging from 60 to
90% depending on which system is used. Table 8-7 summarizes the key features of
each of these technologies. These technologies have been used on boilers burning
low sulfur coal (usually less than 2%) and are attractive alternatives to wet scrub-
bing technology, particularly in the arid western U.S.
8-23
-------
Table 8-7. Key features of dry flue gas desulfurization systems.
Process
Spray dryer with
a baghouse or ESP
Dry injection with a
baghouse or ESP
Combustion of
coal/ limestone
mixture with a
low NO, burner
Sorbents
used
Sodium carbonate
Lime
Limestone
Sodium carbonate
Sodium bicarbonate
Nahcolite
Limestone pellet
Lime
SO,
emission
removal
efficiency
60-90
60-90
75-80
Particulate
emission
removal
efficiency
99 +
99 +
99+*
Development
status
Three utility boilers
(400-500 MW) to be
started up 1981, 1982,
1983. Two industrial
boilers on line.
No commercial installations
planned as of 1980.
EPA currently funding
pilot tests on small
industrial boiler.
*Note: a baghouse or ESP is used for paniculate emission control.
Source: EPA, February 1980.
Spray Dryer with a Baghouse or ESP
The only commercial dry FGD installations in the U.S. at this time use a spray
dryer. Alkaline is injected into a spray dryer with dry particle collection in a
baghouse or ESP. Spray dryers have been used in the chemical, food processing,
and mineral preparation industries over the past 40 years. Spray dryers are vessels
where hot flue gases are contacted with a finely atomized wet alkaline spray. The
high temperatures of the flue gas, 250 to 400°F, evaporate the moisture from the
wet alkaline sprays, leaving a dry powdered product. The dry product is collected
in a baghouse or ESP (Figure 8-7).
Flue gas enters the top of the spray dryer and is swirled by a fixed vane ring to
cause intimate contact with the slurry spray (Figure 8-8). The slurry is atomized
into extremely fine droplets by rotary atomizers. The turbulent mixing of the flue
gas with the fine droplets results in rapid SO2 absorption and evaporation of the
moisture. A small portion of the hot flue gas is added to the spray-dryer-discharge
duct to maintain the temperature of the gas above the dew point. Reheat prevents
condensation and corrosion in the duct. Reheat also prevents bags in the baghouse
from becoming plugged or caked with moist particles.
8-24
-------
Alkaline
Spray dryers
Baghouse
Figure 8-7. Spray dryers with baghouse.
Spray
nozzle
To baghouse
Figure 8-8. Spray dryer.
8-25
-------
Sodium carbonate solutions and lime slurries are the most common absorbents
used. A sodium carbonate solution will generally achieve a higher level of SO2
removal than lime slurries (EPA, February 1980). When sodium carbonate is used,
SO2 removal efficiencies are approximately 75 to 90%, lime removal efficiencies
are 70 to 85% (EPA, February 1980). However, vendors of dry scrubbing systems
claim that their units are capable of achieving 90% SOZ reduction using a lime
slurry in a spray dryer. Lime is very popular for two reasons: lime is less expensive
than sodium carbonate; sodium carbonate and SO2 form sodium sulfite and
sodium sulfate which are very soluble causing leaching problems when landfilled.
Some of the evaporated alkaline spray will fall into the bottom of the spray dryer
and be recycled. The majority of the spray reacts with SO2 in the flue gas to form
powdered sulfates and sulfites. These particles, along with fly ash in the flue gas,
are then collected in a baghouse or electrostatic precipitator (see 413 Student
Manual, EPA 450/2-80-066). Baghouses have an advantage because unreacted
alkaline material collected on the bags can react with any remaining SO2 in the
flue gas. Some process developers have reported SO2 removal on bag surfaces on
the order of 10% (Kaplan and Felsvang, 1979). However, since bags are sensitive to
wetting, a 35 to 50 °F margin above the saturation temperature of the flue gas must
be maintained (EPA, February 1980). ESPs have the advantage of not being as sen-
sitive to moisture as baghouses. However, SO2 removal is not quite as efficient
using ESPs.
The major differences between wet absorption SO2 scrubbers and spray dryer
systems is in the scrubbing method and the amount of moisture during the scrub-
bing action. In a wet scrubber, flue gas is saturated with liquid sprays (usually
alkaline). SO2 is absorbed by the water and also reacts with the chemical. As
previously stated in Chapter 4, absorption increases as temperature decreases.
Flue gas is cooled and saturated with the scrubbing liquid to remove SO2. At
optimum operating temperatures efficiency is usually >90%.
In a spray dryer, finely atomized alkaline droplets are contacted with flue gas
which is at air preheater outlet temperatures (250 to 400 °F). The flue gas is
humidified to within 50 °F of its saturation temperature by the moisture
evaporating from the alkaline slurry. Reaction of the SO2 with the alkaline
material proceeds both during and following the drying process, although to what
degree is not completely understood. Since the flue gas temperature and humidity
are set by air preheater outlet conditions, the amount of moisture that can be
evaporated into the flue gas is also set. This means that the amount of alkaline
slurry that can be evaporated in the dryer is limited by flue gas conditions (see
Appendix I for use of psychrometric chart). Alkaline slurry sprayed into the dryer
must be carefully controlled to avoid moisture in the flue gas from condensing in
the ducting, particulate emission control equipment, or the stack. SO2 removal effi-
ciencies are generally <85% (EPA, February 1980).
8-26
-------
A number of spray dryer systems have been planned for or installed on industrk 1
and utility boilers. These are listed in Table 8-8. Spray dryers will become more
popular as experience with existing units is further documented. They will be par-
ticularly useful in meeting NSPS .regulations for utility boilers burning low sulfur
coal that require only 70% SO2 scrubbing.
Table 8-8. Commercial spray dryer FGD systems.
Station or
plant
Otter Tail Power
Company;
Coyote Station
No. 1, Beulah, ND
Basin Electric;
Antelope Valley
Station No. 1,
Beulah, ND
Basin Electric;
Laramie River
Station No. 3,
Wheatland, WY
Strathmore Paper
Co.;
Woronco, MA
Celanese Corp. ;
Cumberland, MD
Size
(MW)
410
440
500
14
31
Installation
date
6/81
4/82
Spring 1982
12/79
2/80
System
description
Rockwell/ Wheelabrator-
Frye system; four
spray towers in
parallel with 3 atomizers
in each; reverse air —
shaker baghouse with
dacron bags
Joy/Niro Atomizer Inc.;
five spray modules with
single rotary atomizer
in each; reverse air
baghouse with Teflon
coated fiberglass bags
Babcock and Wilcox;
four spray reactors with
12 "Y-jet" nozzles
in each; electrostatic
precipitator
Mikropul; spray dryer
and pulse jet
baghouse
Rockwell/Wheelabrator-
Frye; one spray tower
followed by a baghouse
Sorbent
Soda ash
(sodium
carbonate)
Lime
Lime
Lime
Lime
Coal sulfur
content
(%)
0.78
0.68-1.22
0.54-0.81
2-2.5
1-2
SO, emission
removal
efficiency
<%)
70
62-78
85-90
75
85
Source: EPA, February 1980.
Dry Injection
In dry injection systems, a dry alkaline material is injected into a flue gas stream.
This is accomplished by pneumatically injecting the dry sorbent into a flue gas
duct, or by precoating or continuously feeding sorbent onto a fabric filter surface.
Most dry injection systems use pneumatic injection of dry alkaline material in the
boiler furnace area or in the duct that precedes the ESP or baghouse. Sodium-
based sorbents are used more frequently than lime. Many dry injection systems
have used nahcolite, a naturally occurring mineral which is 80% sodium bicar-
bonate found in large reserves in Colorado. Sodium carbonate (soda ash) is also
used but is not as reactive as sodium bicarbonate (EPA, February 1980). The
major problem of using nahcolite is that it is not presently being mined on a com-
mercial scale. Large investments must be made before it will be mined commer-
cially. Other natural minerals such as raw trona have been tested; trona contains
sodium bicarbonate and sodium carbonate.
8-27
-------
Dry injection systems have been tested at a number of power stations throughou,;
the U.S. Descriptions of these pilot systems can be found in Survey of Dry SO2 Con-
trol Systems (EPA, February 1980). The major problems with dry injection systems
are the low sodium utilization in the process and the disposal of leachable sodium-
sulfur compounds. EPA reports that only 40 to 60% of the dry alkaline injected
material is used at high SO2 removal conditions (EPA, February 1980).
Other Dry SOt Processes
Coal and limestone fuel mixtures have been tested as a method for reducing SO2
emissions. This technology using limestome pellets in a fluidized bed combustion
boiler or limestone injection with a low NO* burner was discussed earlier in this
chapter. Another promising dry process, the Shell UOP process, uses a copper
oxide catalyst for SO2 emission reduction. This process was discussed in detail in
Chapter 7 of this manual.
Comparison of FGD Systems
Comparisons between dry and wet FGD systems can be made in four major areas:
waste disposal, chemical reagent requirements, SO2 removal efficiencies, and
economics. A comparison of the FGD systems covered in this manual will follow.
Lime, limestone, double alkali, and dry FGD systems produce a sludge that must
be disposed of properly. New installations must meet all solid waste regulations
including stringent RCRA regulations. The regenerable FGD processes —including
Wellman-Lord, magnesium oxide, and citrate —generate a usable product (sulfur
or sulfuric acid) that has a commercial value. The dry FGD systems produce a dry
waste product that can be discarded using conventional fly ash removal systems
instead of sludge removal systems. However, sodium-based dry FGD systems are
undesirable because they produce waste products that easily leach from conven-
tional landfills.
Lime and limestone are cheaper than sodium based absorbents and are readily
available. In dry scrubbing systems, a higher stoichiometric ratio of lime is
necessary to achieve SO2 reduction efficiencies similar to wet lime and limestone
scrubbing efficiencies. Consequently, costs for lime or limestone materials will be
somewhat higher in dry systems.
Regarding SO2 removal efficiencies, all of the wet scrubbing processes are
capable of at least 90% removal efficiency. Most units are capable of 95% removal
efficiency. These systems have been installed and operating for a number of years
with good reliability. Dry scrubbing is capable of removing at least 75 to 85% SO2.
Vendors of dry FGD equipment claim 90% capabilities. Dry FGD systems are
capable of meeting the NSPS regulation (1978) of 70% scrubbing when SO2 emis-
sions are less than 0.6 lb/106 Btu. This technology is very attractive for utility
boilers burning low sulfur western coal. With improvements, dry scrubbing may be
used on boilers burning medium sulfur coal (2 to 3%) in the near future.
8-28
-------
Cost Comparisons
Cost estimates for FGD systems have been made by various groups including the
EPA, Tennessee Valley Authority (TVA), and the Electric Power Research Institute
(EPRI). Two articles in the Journal of the Air Pollution Control Association
(JAPCA, April and May 1981) summarize FGD systems on an environmental and
cost basis. Cost estimates of selected FGD systems are given in Table 8-9. This table
was constructed using cost estimate data from a paper presented at the FGD sym-
posium in Houston, Texas, October, 1980 (McGlamery et al., 1980). The paper
presented cost estimates of FGD systems based on 1979 capital costs and 1980
operating costs. The authors of this paper also worked up a cost estimate of a
limestone scrubber based on 1982 capital costs and 1984 operating costs. As can be
seen from Table 8-9, the cost of a 500 MW power plant burning 3.5% sulfur coal,
meeting the 1.2 Ib SO2/106 Btu NSPS requirement, was 98 $/kW for capital cost
and 4.02 mills/kWh operating costs.
Table 8-9. Cost estimates of selected FGD systems.*
FGD system
Limestone
Lime
Double alkali
Wellman-Lord
Magnesium oxide
Citrate
Mid-1979 capital
investment
($/kW)
98
90
101
131-141
132
143
Mid-1980 first year
operating cost
(mills/kWh)
4.02
4.25
4.19
5.11-6.03
5.08
6.44
*Cost basis: 500 MW power plant burning 3.5% sulfur coal achieving
90% SO2 removal efficiency.
Source: McGlamery, 1980.
The costs of a limestone scrubber on a similar power plant with the following
new premises was also calculated:
• meeting 1978 NSPS requirements
• spare absorber and emergency bypass
• ID booster fans instead of FD fans
• ponding and landfill estimates to meet RCRA requirements
• forced oxidation of sludge
• use of a spray tower instead of a mobile bed scrubber
• changes in economic premise including 1982 capital investment and 1984
operating costs
The cost of a limestone scrubber under this new premise is 194 $/kW capital
cost and 9.9 mills/kWh operating costs. These cost estimates are almost double
those given in Table 8-9. Therefore, the cost data given in Table 8-9 should be
used for comparison of FGD systems only.
8-29
-------
Over the past 15 years, a wealth of material has been written and documented
concerning FGD control technology. The authors of this manual suggest that the
readers turn to the many publications from EPA-IERL concerning this subject,
particularly the proceedings from the annual FGD symposiums sponsored by the
EPA.
References
Banks, R. R., and Hochhauser, M. L. 1978. Desulfurization of Flue Gas with
Nahcolite in a Pilot Fabric Filter. Air Poll. Control Assoc. Meeting Paper,
Houston, Texas, June 25-30, 1978: 78-46.4.
Beachler, D. S. 1978. California Air Resources Board Memorandum: Concerning
the Relationship Between Total Suspended Paniculate, Nitrogen Oxide, Sulfur
Oxide and Hydrocarbon Emissions in the South Coast Air Basin.
Costle, D. M. 1979. New Source Performance Standards for Coal Fired Power
Plants./ of Air. Poll. Control Assoc. 29:690-692.
Crowe, R. B. 1981. First Year Operational Experience Celanese Fibers Company
Coal-Fired Boiler Using a Dry Flue Gas Cleaning System. Air Poll. Control
Assoc. Meeting Paper, Philadelphia, June 21-26, 1981: 81-35.2.
Environmental Protection Agency (EPA, 1981). Proceedings: Symposium on Flue
Gas Desulfurization—Houston, Texas, October 1980, Volumes I and II. EPA
600/9-81-019a and EPA 600/9-81-019b.
Environmental Protection Agency (EPA). August, 1980. Research Summary, Con-
trolling Sulfur Oxides. EPA 600/8-80-029.
Environmental Protection Agency (EPA). February, 1980. Survey of Dry SO2 Con-
trol Systems. EPA 600/7-80-030.
Environmental Protection Agency (EPA). November, 1979. Decision Series, Sulfur
Oxides Control in Japan. EPA 600/9-79-043.
Environmental Protection Agency (EPA). July, 1979. Proceedings: Symposium on
Flue Gas Desulfurization—Las Vegas, Nevada, March, 1979, Volumes I and II.
EPA 600/7-79-167a and EPA 600/7-79-167b.
Environmental Protection Agency (EPA). May, 1979. Decision Series, Sulfur
Emission: Control Technology and Waste Management. EPA 600/9-79-019.
Environmental Protection Agency (EPA). February, 1979. Summary Report-
Sulfur Oxides Control Technology Series: Flue Gas Desulfurization, Wellman-
Lord Process. EPA 625/8-79-001.
Environmental Protection Agency (EPA). January, 1979. Evaluation of Dry
Sorbents and Fabric Filtration for FGD. EPA 600/7-79-005.
8-30
-------
Environmental Protection Agency (EPA). January, 1979. Electric Utility Steam
Generating Units—Flue Gas Desulfurization Capabilities as of October 1978.
EPA 450/3-79-001.
Environmental Protection Agency (EPA). March, 1978. Flue Gas Desulfurization
System Capabilities for Coal-Fired Steam Generators, Volume II, Technical
Report. EPA 600/7-78-032b.
Environmental Protection Agency (EPA). March, 1978. Proceedings: Symposium
on Flue Gas Desulfurization—Hollywood, FL, November 1977, Volumes I and
II. EPA 600/7-78-058a and EPA 600/7-78-058b.
Environmental Protection Agency (EPA). March, 1978. Flue Gas Desulfurization
Systems: Design and Operating Considerations, Volume I, Executive Summary.
EPA 600/7-78-030a.
Environmental Protection Agency (EPA). March, 1978. The Effect of Flue Gas
Desulfurization Availability on Electric Utilities, Volume I, Executive Summary.
EPA 600/7-78-031a.
Environmental Protection Agency (EPA). March, 1978. Flue Gas Desulfurization
System Capabilities for Coal-Fired Steam Generators, Volume I, Executive Sum-
mary. EPA 600/7-78-032a.
Environmental Protection Agency (EPA). March, 1978. The Effect of Flue Gas
Desulfurization Availability on Electric Utilities, Volume II, Technical Report.
EPA 600/7-78-031b.
Environmental Protection Agency (EPA). November, 1977. Flue Gas Desulfuri-
zation System Manufacturers Survey. EPA 450/3-78-043.
Environmental Protection Agency (EPA). 1977. EPA Technology Transfer Capsule
Report, First Progress Report. Wellman-Lord SO2 Recovery Process—Flue Gas
Desulfurization Plant. EPA 625/2-77-011.
Environmental Protection Agency (EPA). EPA Technology Transfer Capsule
Report, Third Progress Report. Lime/Limestone Wet-Scrubbing Test Results at
the EPA Alkali Scrubbing Test Facility. EPA 625/2-76-010.
Environmental Protection Agency (EPA). EPA Technology Transfer Capsule
Report, Second Progress Report. Lime /Limestone Wet-Scrubbing Test Results at
the EPA Alkali Scrubbing Test Facility. EPA 625/2-75-008.
Gentile, A. R. and Dunkle, D. M. 1979. Operation of Commercial FGD Sludge/
Fly Ash Waste Treatment Systems. Air Poll. Control Assoc. Meeting Paper, Cin-
. cinnati, Ohio, June 24-29, 1979: 79-31.6.
Hatfield et al. 1979. Six Years' Operation and Maintenance Experience with a
Wellman-Lord SO2 Abatement System at a Sulfuric Acid Plant. Air Poll. Con-
trol Assoc. Meeting Paper, Cincinnati, Ohio, June 24-29, 1979: 79-23.3.
Hollett, G. T. 1979. Dry Removal of SO2-Application to Industrial Coal-Fired
Boilers. Air Poll. Control Assoc. Meeting Paper, Cincinnati, Ohio, June 24-29,
1979: 79-23.1.
8-31
-------
Jahnig, C. E. and Shaw, H. 1981. A Comparative Assessment of Flue Gas
Treatment Processes Part I - Environmental and Cost Comparison./, of Air Poll
Control Assoc. 31:421-428.
Jahnig, C. E. and Shaw, H. 1981. A Comparative Assessment of Flue Gas
Treatment Processes Part II - Environmental and Cost Comparison./ of Air
Poll. Control Assoc. 31:596-604.
Jones, D. G. et al. 1979. Lime/Limestone Scrubber Operation and Control. Air
Poll. Control Assoc. Meeting Paper, Cincinnati. Ohio, June 24-29, 1979: 79-23.6
Kaplan, S. M. and Felsvang, K. 1979. Spray Dryer Absorption of SO* from
Industrial Boiler Flue Gas. 86th National AICHE Meeting Paper, Houston,
Texas, April, 1979.
Link, W. F. and Mann, E. L. 1979. Operating Experience with the Wellman-
Lord/Allied Chemical Flue Gas Desulfurization Plant at Northern Indiana
Public Service Company's Dean H. Mitchell Station. Air Poll. Control Assoc.
Meeting Paper, Cincinnati, Ohio, June 24-29, 1979: 79-31.7.
Madenburg et al. 1979. Citrate Process Demonstration Plant-Startup and
Operation. Air Poll. Control Assoc. Meeting Paper, Cincinnati, Ohio Tune
24-29, 1979: 79-23.4.
McGlamery et al. 1980. FGD Economics in 1980. Flue Gas Symposium Paper,
Houston, Texas, October 28-31, 1980.
Milliken, J. O. and Young, C. W. 1981. Emissions Performance of Georgetown
University Fluidized Bed Boiler. Air Poll. Control Assoc. Meeting Paper,
Philadelphia, PA, June 21-26, 1981: 81-35.3.
Mobley, J. D. and Lim, K. J. 1981. Control of Gaseous Pollutants from Indus-
trial Combustion by Chemical Reaction. Air Poll. Control Assoc. Meeting Paper,
Philadelphia, PA, June 21-26, 1981: 81-35.7.
Molburg, J. 1980. A Graphical Representation of the New NSPS for Sulfur
Dioxide./, of Air Poll Control Assoc. 30:172.
Murphy, J. T. 1979. Start-Up Experience with Lime Scrubbing at Pleasants
Power Station. Air Poll. Control Assoc. Meeting Paper, Cincinnati, Ohio June
24-29, 1979: 79-31.4.
Reiners, M. H. et al. 1981. Operating Experience with the Largest High Sulfur
FGD System. Air Poll. Control Assoc. Meeting Paper, Philadelphia, PA, June
21-26, 1981: 81-35.1.
Richman, M. 1979. FGD System Operation at Martin Lake Steam Electric Station.
Air Poll. Control Assoc. Meeting Paper, Cincinnati, Ohio, June 24-29 1979-
79-31.5.
8-32
-------
Spencer, H. W. et al. 1981. Experience with Baghouses for Dry FGD Service.
Air Poll. Control Assoc. Meeting Paper, Philadelphia, PA, June 21-26 1981-
81-9.5.
Trijonis, J. 1978. Visibility in the Southwest —An Exploration of the Historical
Data Base. Air Poll. Control Assoc. Meeting Paper, Houston, Texas, June 25-30
1978: 78-43.7.
8-33
-------
Chapter 9
Industrial Exhaust Systems
Hoods
Control of all air pollutants begins at their point of capture, the hood. The overall
control efficiency that can be achieved is a function of the capture efficiency of the
hood. For example, if a poorly designed hood captures only 70% of the emissions
generated, then the overall efficiency of the system can never be greater than 70%
even if the control device is 100% efficient. Any pollutant can be captured effec-
tively with an appropriately sized hood and a fan with enough power to supply the
suction. Hoods are designed, however, to achieve the maximum pollutant capture
efficiency with a minimum of air flow and power consumption.
Hoods are generally classified according to their shape and the manner in which
they capture the contaminant. Hoods are referred to as enclosure or nonenclosure
depending on the manner in which they capture the pollutant. Some common
hood shapes are the slot, plain opening, booth, and canopy (Figure 9-1). Hood
entry coefficients will be discussed later in this chapter. Enclosure hoods may
totally or only partially enclose the source of emissions. The term "enclosure" is
somewhat of a misnomer since some of these hoods can be located close to the
point of emissions so that capture occurs before dispersion of the contaminant can
take place. For example, in many paint spraying operations a booth is used to cap-
ture the emissions generated. Occasionally the painting is done inside the booth or
it may be done a few feet in front of the booth. Both of these operations would be
an example of an enclosure hood.
Nonenclosure hoods refer to hoods that are usually placed around the perimeter
of the contaminant-generating source. The most common nonenclosure hood used
to capture gaseous emissions is the lip or rim exhaust located on open surface
tanks.
9-1
-------
Figure 9-1. Typical hoods and their entry coefficients.
Hood Design
The air volume or suction required for a hood depends on: hood shape, hood type,
hood size, required capture velocity, distance of the hood from the source of
pollutants, and temperature of the contaminant exhaust stream. Due to the
numerous variables, equations for predicting the required exhaust volume are
based on empirical data for a particular type of hood.
9-2
-------
The basis for hood design equations starts with the simplest case —air flowing
into a circular hood or duct. From studies done by Dalle Valle and others, the
actual flow pattern of air being sucked into the hood is illustrated in Figure 9-2.
The curved radial lines in Figure 9-2 are referred to as contour lines and represent
lines of constant velocity. The lines perpendicular to the duct are streamlines and
represent the direction of flow.
o
tti
^
(A
"bo
O.
o
o
Constant
velocity line
Streamline
Air flow
Source: Industrial Ventilation, 1976.
Figure 9-2. Flow diagram of air being pulled into a plain opening.
Two important generalizations are illustrated in Figure 9-2. First, a rapid
decrease in velocity occurs with increased distance from the hood face. Therefore,
the closer the hood is to the source (or by totally enclosing the source), the more
efficient and effective the capture of pollutants. Secondly, air is drawn into the
openings from all directions, meaning that more air is drawn in than required.
Flanges can be used to eliminate air pull from areas where no contaminants are
present. Figure 9-3 illustrates the effect adding flanges has on the contour and
streamlines. For most cases it is estimated that flanges can reduce air requirements
by 25% and need never extend more than 6 inches beyond the hood entrance
(Industrial Ventilation, 1976).
9-3
-------
Flange
Constant
velocity line
Streamline
Air flow
Source: Industrial Ventilation, 1976.
Figure 9-3. Effect of flange on air flow into a plain opening.
The following design equations for hood systems are categorized according to the
shape of hood. All equations are for cold contaminant exhaust streams unless
otherwise stated. A more extensive review of hood types and their design equations
is given in Industrial Ventilation (1976).
Total Enclosure Hood
For any totally enclosed hood or for emissions generated inside a booth, the air
flow into the hood can be computed by:
(Eq. 9-1)
Q,=
Where: Q= air flow (into hood face), cfm
v=air velocity required to capture the pollutant, fpm
A = total open area of the hood, ft2*
The capture velocity is the velocity necessary (at any point) to overcome opposing
air currents and to capture the air contaminants. Table 9-la presents a range of
capture velocities for certain processes (Industrial Ventilation, 1976).
*The units in this chapter are expressed in English rather than metric, since most published
tables and charts are in English units.
9-4
-------
Table 9-la. Range of capture velocities and duct velocities.
Condition of dispersion
of contaminant
Examples
Capture velocity
(fpm)
Released with practically
no velocity into quiet air
Evaporation from tanks;
degreasing, etc.
50-100
Released at low velocity
into moderately still air
Spray booths; intermittent
container filling; low speed
conveyor transfers;
welding; plating; pickling
100-200
Active generation into
zone of rapid air motion
Spray painting in shallow
booths; barrel filling;
conveyor loading; crushers
200-500
Released at high initial
velocity into zone of very
rapid air motion
Grinding; abrasive
blasting, tumbling
500-2000
In each category above, a range of capture velocity is shown. The proper choice
of values depends on several factors:
Lower end of range
1. Room air currents minimal or
favorable for capture.
2. Contaminants of low toxicity or of
nuisance value only.
3. Intermittent, low production.
4. Large hood —large air mass in
motion.
Upper end of range
1. Disturbing room air currents.
2, Contaminants of high toxicity.
3. High production, heavy use.
4. Small hood —local control only.
Source: Permission granted by Committee on Industrial Ventilation, American
Conference of Governmental Industrial Hygienists, Box 16153, Lansing, MI
48901. Taken from 14th Edition (1976) of Industrial Ventilation.
Table 9-lb. Recommended minimum duct velocities.
Nature of contaminant
Gases, vapors, smokes,
fumes, and very light
dusts
Medium-density dry dust
Average industrial dust
Heavy dusts
Examples
All vapors, gases, and
smokes; zinc and
aluminum oxide fumes;
wood, flour, and cotton
lint
Buffing lint; sawdust;
grain, rubber, and plastic
dust
Sandblast and grinding
dust, wood shavings,
cement dust
Lead and foundry shake-
out dusts; metal turnings
Duct velocity
(fpm)
2000
3000
4000
5000
Source: EPA, 1973.
9-5
-------
Free Standing Plain Hood
For a free standing or unbounded hood, Equation 9-1 must be modified to account
for the distance a source is from the hood. Dalle Valle developed Equation 9-2 for
air flowing into a plain, circular duct or a rectangular hood that is essentially
square.
(Eq. 9-2) Q.= vx(10*2 + A)
Where: Q= air flow, cfrn
v* = air velocity required for capture at some distance (x) from hood
opening, fpm
x - distance of contaminant source from the center of the hood face, ft
A = area of hood face, ftz
Equation 9-2 was developed from Figure 9-2 and it also shows the rapid increase
in air volume required for capture (due to decrease in velocity) as the source gets
farther from the hood. The volume increases in relation to the distance squared.
Equation 9-2 applies to a free standing or unobstructed hood. For a square or
rectangular hood that is bounded on one side by a flat surface (the floor or wall)
Equation 9-2 then becomes:
(Eq.9-3) Q=v
Example 9-1
Determine the total volume of air entering a paint spray booth that is 10 ft wide
and 7 ft high. Some objects are painted outside the booth but never at a distance
of more than 3 ft. Room air currents are kept low.
Solution:
The paint spray booth would be a plain opening bounded on one side by a flat sur-
face (the floor). Therefore, Equation 9-3 applies.
The capture velocity can be estimated from Table 9- la. For paint spray emitted
into a draftless area the capture velocity should be 100 fpm.
Substituting into Equation 9-3:
= 100
= ll,500cfm
9-6
-------
An important point to note is that in order to capture the pollutant the velocity at
3 ft from the hood face must be at least 100 fpm. If painting was done farther
away from the hood face, the velocity still must be the same. The volume of air
would then greatly increase.
For example, if painting is done at 5 ft instead of 3 ft from the hood, then this
raises the required air volume pulled into the system to 19,500 cfm.
A common way of expressing control volume is based on the square feet of hood
opening. For example:
Q 11,500 cfm
A (7 ft)(10 ft)
164.3 cfm of air
sq ft of hood opening
Canopy Hood: Cold Processes
Canopy hoods have three or four open sides and are almost identical to plain free-
standing hoods. The major difference is that canopies are used over the top of
open tanks. The canopy portion of the hood is normally extended beyond the edge
of the tank to minimize any side draft disturbances. The air volume for an open
canopy hood capturing emissions from a cold process is:
(Eq. 9-4) Q=lAhplv
Where: Q= air flow of an open canopy hood, cfm
hp = canopy perimeter excluding closed sides, ft
/ = vertical distance between canopy and tank hood, ft
v = control velocity, fpm
Canopy Hood: Hot Processes
The design of hoods to capture gaseous emissions from a hot process is markedly
different from that for a cold process. Expansion of the hot gases due to the ther-
mal effect must be taken into account. As the hot gas rises, it mixes turbulently
with the surrounding air. The higher the gas column rises, the larger it becomes,
and the more it is diluted with ambient air. Equation 9-5 is an empirical relation-
ship developed by Hemeon (1955) for predicting the expansion of the rising plume
for high canopy hoods (located more than 3 ft above the heat source):
(Eq. 9-5) DC = 0.5XC088
Where: Dc = diameter of the hot column of air at the level of the hood face, ft
Xc = distance from the hypothetical point source to the food face, ft
9-7
-------
Xc is defined in Figure 9-4 as the sum of the distance from the hot source to the
hood face, Yc, plus the distance from the top surface of the heat source to a
hypothetical vortex of the fume column, Z.
«s*?r*7
Imaginary point
source
Source: Hemeon, 1955.
Figure 9-4. Dimensions used in designing high canopy hoods
controlling emissions from hot sources.
Values of Zc can be found from:
(Eq. 9-6)
C = (2DS)1
Where: Ds = width of the hot source
Zc = distance from top of hot source to hypothetical fume vortex, ft
9-8
-------
The actual diameter of the hood opening Dh must extend beyond the rising column
of hot gases to ensure complete capture. Most designers increase the hood diameter
by a factor of 0.8 Yc (Industrial Ventilation, 1976). Therefore, diameter of the
canopy is given by:
(Eq. 9-7) Dfc = Dc + 0.8Yc
The volume of gas entering the canopy hood includes some ambient air and is
given by:
(Eq. 9-8) Q.= vAc + vr(A - A.)
Where: Q= total volume of air entering the hood, cfm
v = velocity of the rising gases at the hood face, fpm
Ac = area of the rising column of gas at the hood face, ft2
vr = required velocity through the remaining area of the hood,
A- AC, fpm
A = total area of hood face, ft2
The value of vr depends on the draft in the hood area, height of the hood above
the source, and degree of capture required. Values of vr usually range between 100
and 200 fpm, with higher values giving higher pollutant capture efficiencies.
The control velocity for capturing emissions from a flat, horizontal heat source is
given by:
(Eq.9-9) y
Where: v= control velocity, fpm
A,= top surface area of hot source, ft2
AT = temperature difference between the hot source and ambient, °F
Xc = distance from hood opening to imaginary point source, ft
Control velocities for other, irregularly shaped hot sources can be found in
Cheremisinoff (1976).
9-9
-------
It is important to note that Equations 9-6 through 9-9 are for high circular
canopy hoods controlling emissions from a flat, horizontal hot source. Equations
for low canopy hoods are listed in Table 9-2.
Table 9-2. Control volumes for low canopy hoods.
Shape of heat source
Circular
Rectangular
Control volume
Q.= 4.7(D,,)233(AT)04J
Q.= 6.2(w./,)'M(AT)a«
Where: D, = diameter of hood, ft
/* = length of hood, ft
w = width of hood, ft
AT = difference between the temperature of the heat
source and ambient, °F
Example 9-2
A zinc-melting pot 5 ft in diameter contains molten metal at 800 °F. A high canopy
hood is used to capture emissions and is located 8 ft above the pot. Assuming
ambient air temperature to be 80 °F and using vr as 200 fpm, calculate the hood
size and exhaust rate.
Solution:
1. To calculate the hood size, first find the distance from the top of the pot sur-
face down to the hypothetical point source, Z0 from Equation 9-7.
= (2X5 ft)1138
= 13. 7 ft
The total height Xc is given by:
= 13.7 + 8
= 21.7 ft
The diameter of the rising column of gases at the hood face is given by
Equation 9-5.
DC = 0.5XC°88
= 0.5(21. 7)° 88
= 7.5 ft
The hood size required, including increase for the rising hot gas is given
by Equation 9-7.
D_ = DC + 0.8YC
= 7.5 ft + 0.8(8)
= 7.5 ft + 6. 4 ft
= 13. 9 ft
9-10
-------
2. To calculate the hood exhaust rate, use Equation 9-8.
To use Equation 9-8, first compute the area of the melting pot and the velocity
of rising gas at hood face.
Area of hot source:
= f(5)!
= 19.6 ft
The velocity is computed from Equation 9-9.
_ 8(A.)0-33(AT)0-"
Xc°25
_ 8(19. 6)° "(800- 80)° «
(21.7)02*
= 157 fpm
The area of the rising column of gas is:
= 44.2 ft2
The area of the hood face is:
A'T-v
= —(13.9)2
4
= 151.7 ft2
Substituting into Equation 9-8 for the total volume of air:
= 157(44.2) + 200(151.7 - 44.2)
= 28,440 cfm
9-11
-------
Nonenclosure Lip Hood
One of the most common types of nonenclosure hoods for controlling vapors from
open surface tanks is the lip hood pictured in Figure 9-5 (these hoods are
sometimes called slot hoods). These hoods are rarely used with hot gases. The lips
are designed for a velocity of approximately 200 fpm through the lip face at the
required ventilation rate (EPA, 1973).
Lip exhausts
(slots)
Figure 9-5. Open top tank degreaser with lip exhausts.
For a tank with two parallel lip hoods, the ventilation rate required and the lip
width may be estimated from Figure 9-6. If a lip hood is used on only one side of a
tank and the opposite side of the tank is bounded by a vertical wall, Figure 9-6 can
be used to estimate the ventilation rate. The procedure, as outlined by the EPA
(1973), is to double the tank width and read a ventilation rate from Figure 9-6.
The actual ventilation rate is estimated as one-half the value read from the graph.
The following example illustrates the use of Figure 9-6.
9-12
------- An error occurred while trying to OCR this image.
-------
Hood Entry Coefficients
When air flows into a hood or duct, pressure losses occur which result in a
decreased flow rate. These losses are due to turbulence and are dependent on the
shape of the hood or duct. The coefficient of entry (C.) indicates the extent of
these losses. For example, in a theoretically perfect hood with no turbulence loss,
C,= 1.0. Figure 9-1 lists the entry coefficients for some typical hoods. The coeffi-
cient of entry is used to determine the actual flow rate and pressure after the air
enters the hood.
Ducts
Air Flow in Ductwork
In describing air flow through a system, a number of pressure terms are used. The
three most important are: static pressure, velocity pressure, and total pressure. Air
flowing in a pipe is acted on by static and velocity pressure, which when added
together give the total pressure (Figure 9-7).
Static pressure
0.25 in.
Velocity pressure
1.0 in.
Total pressure
1.25 in.
Figure 9-7. Pressure terms used to describe air flow in a duct (measured in inches).
Any gas within a confined enclosure will have a static pressure, whether or not
the gas is in motion. Static pressure, p,, acts in all directions and is independent of
the velocity of the gas. Therefore, static pressure is measured at right angles to the
direction of air flow to avoid influence from the air velocity. In duct design, the
friction losses in a pipe can sometimes be measured by the difference in static
pressure (no change in direction or velocity). Therefore, the static pressure is
sometimes called the frictional or resistance pressure. Static pressure can be positive
or negative, compared to the local atmosphere.
The velocity pressure, pv, is the pressure created by air traveling at a specific
velocity- Velocity pressure acts onlyin th&jdireGtion-nf flow therefore it is always
_3osmve_The velocity pressure is obtained by measuring the total pressure and sub-
tracting out the static pressure. There is a basic relationship between the velocity of
-v
9-14
-------
a gas and the velocity pressure. For an air stream with a
(standard conditions) the air velocity can be computed from the velocity pressure
using the following equation (Industrial Ventilation, 1976):
(Eq. 9-10) v = 4005VpT
Where: y=air velocity, fpm
pv = velocity pressure, in. H2O
Equation 9-10 is derived by applying Bernoulli's equation for air flow in a duct.
(The derivation of the equation is given in APTI Course 450 Student Manual,
Appendix C.) Equation 9-10 is used extensively in stack sampling, where it is com-
monly known as the pilot tube equation. The common form of the pilot tube
equation is:
(Eq. 9-11)
Where: v = gas velocity, ft/sec
T,= temperature of the stack, °R
M, = molecular weight of the stack gas, Ib/lb mol
P,= absolute pressure of the stack, in. Hg
Ap = velocity pressure, in. H2O
The above equation equals Equation 9-10 when T,= 530°R, Ms= 29 Ib/lb mol and
P, = 29.92 in. Hg. Note that in stack sampling the velocity pressure is symbolized by
Ap, while in industrial ventilation it is pv.
The sum of the static and the velocity pressure at a given point in the duct is the
total pressure, PT (Figure 9-7). Total pressure is measured in the direction of the
flow. Therefore, the total pressure is sometimes referred to as the impact or
dynamic pressure. The total pressure should not be confused with the absolute
pressure. The absolute pressure is the_surn_of the_a£mospheric pressure plus^the,
r_pressurejof the system (Chapter 2). The system pressure can be either positive or
negative.
Air Flow Resistances or Losses in a Duct
Pressure losses from air flowing in a pipe are due to friction and turbulence. Static
or friction pressure losses are caused by air abrading or rubbing against the sides of
the duct. Turbulent losses are due to rapid changes in direction or velocity. A
change in duct cross sectional area will change the velocity of the gas in the duct.
The sum of the friction and turbulent losses over a specific length of pipe is termed
the pressure drop. The pressure drop of a system is determined by measuring the
difference in total pressure at two points in the system.
9-15
-------
The friction and dynamic losses can be further subdivided into five categories
These losses are:
Energy required to accelerate a volume of air from rest. Inertia
losses are essentially the same as the velocity pressure.
Orificejosses^ Pressure losses at the hood or duct entrance due to turbulence
Orifice losses are dependent on the shape of the opening and are measured by the
entry coefficient, C..
Straight duc^friction losses: Pressure losses due to air rubbing along the sides of
the-TiucrrlvIiny-charts and tables have been developed that give the friction losses
m straight ducts. Figure 9-8 plots the friction losses in inches of water per 100 feet
of duct versus air volume with duct diameter and velocity as parameters. If any two
of the above quantities are known, the other two can be read from the chart.
Elbow and_brandijntryjosses: Pressure losses due to change in the streamlines
of^fTairTtream moving through an elbow or entering a branched connection
These losses are computed as equivalent feet of straight duct that will have the
same pressure loss. Equivalent lengths of elbows and entries are given in Table 9-3
(Industrial Ventilation, 1976).
For gradual contractions in the cross sectional
losses are small (Alden, 1970). Abrupt contractions are
rare in well-designed exhaust systems except as outlets from chambers. For duct
expansions, abrupt enlargements result in a greater pressure loss.
Checking the Design of an Exhaust System
A number of methods and general rules of practice are used in the initial design of
an exhaust system. These procedures are covered in detail in Industrial Ventilation
(1976). This section will concentrate on a procedure to check the flows and
pressures of an existing exhaust system. The objectives of checking an exhaust
system are (EPA, 1973):
1 . to determine the exhaust volume and suction velocity at each point and
evaluate the adequacy of contaminant pickup.
2. to determine the system's static and total pressure; and evaluate fan capacity,
speed, and horsepower.
3. to determine the total exhaust volume and evaluate the size and performance
of the control device.
If all calculations are put in tabular form, a systematic arid orderly procedure
can be followed, whether designing a new system or checking the design of an
existing system.
The following illustrative problem is based on a procedure outlined by Alden
(1970) for duct design. Flow calculations take place in two steps. The first, Table
9-4a, establishes the minimum pressure in the main section of the duct that will
induce the desired flows in each branch. This is done by computing the pressure at
each junction of the main at a branch. This pressure includes both hood entry and
friction losses in the branch. These values in Table 9-4a are only trial figures and
are used as the starting point for the second set of calculations, Table 9-4b. The
second set of calculations determines the proper pressures and flow rates.
9-16
------- An error occurred while trying to OCR this image.
-------
Solution:
The following notes refer to the column and line numbers in Table 9-4a.
Column 1
Column 2
Column 3
Column 4
Column 5
Column 6
Column 7
Column 8
Column 9
Column 10
Column 11
Column 12
Column 13
Column 14
Self explanatory.
Self explanatory.
Required exhaust volume pickup at each hood. These values are given in the
or°s?lT Stfa"ment- If 'hey are not suPPlie<*. values can be obtained from Table 9-1
or similar tables in other references.
Obtained from sketch of plant, Figure 9-9.
Determined by dividing the volumetric flow (Column 3) by the area of the duct Note
that all velocities must be greater than 2000 fpm (from Table 9-1)
Obtained from the sketch in Figure 9-9.
Equivalent length is obtained from Table 9-3 for the given elbow angles
1 from Table 9-3 for the ^ branch '
Read from Figure 9-8, the friction chart
by lo ""^ by
Obtained from Figure 9-1.
Negative pressure at the hood required to produce the exhaust volume (Column 3) at
this given set of conditions. Column 13 is computed using Equation 9-10
Sum of Columns 11 and 13. This is the minimum negative suction pressure at the
junction of the mam, which will produce the proposed flow in the branch
9-20
------- An error occurred while trying to OCR this image.
-------
The following notes refer to the column and line numbers in Table 9-4b.
Column 15 Self explanatory.
Column 16 Self explanatory.
Column 17 Transferred from Column 14 of Table 9-4a.
Column 18
Line 1 Pressure drop across branch BC. Calculation starts at the hood for which the pressure
drop from hood to fan inlet is the greatest (usually this is the farthest hood from the
fan).
Column 18
Line 2 Pressure drop across section AC which joins branch BC. Since the pressure drop from
the junction at C to the atmosphere at A and B must be equal for both branches, the
junction pressure must be the same.
Column 19
Line 2 Computes the effect of the hood suction at A due to the increased pressure at
junction C. Hood suction increases proportionally to an increase in junction pressure
(Column 18/Column 17).
Column 20
Line 2 Gives the effect of the high pressure on the branch volume. Branch volume is related
to pressure by:
Corrected cfm = design rfm| /Pressure of branch with larger Ap
V Pressure of branch with smaller Af
Ap
Column 21
Line 2
Column 22
Line 2
Corrected branch volumetric flow rate.
Expected hood suction.
The next section of columns examine flow through the main:
Column 23
Line 3
Column 24
Column 25
Column 26
Column 27
Column 28
Column 29
Column 30
Column 31
Sum of the corrected branch volumes entering the main at C (AC plus BC).
Read from the friction chart in Figure 9-8.
Taken from the sketch, Figure 9-9.
Taken from the sketch of the system, Figure 9-9.
Given.
Read from Table 9-3.
Sum of Columns 26 and 28.
Read from Figure 9-8 for the conditions given in Columns 23 through 25.
Product of Columns 29 and 30, divided by 100. This number is added to Column 17,
line 4 and entered in Column 18, line 4. This is due to the fact that the pressure in
Column 17, line 4 is greater than Column 18, line 2.
The total static pressure supplied by the fan is computed by adding lines 6, 7
and 8 of Column 18. The computations proceed in a similar manner, with the
negative pressure of the fan inlet found to be 3.28 in. H2O. The adsorber bed
pressure is entered in Column 18, line 8 and the total pressure that must be
developed by the fan is 6.76 in. H2O.
9-22
------- An error occurred while trying to OCR this image.
-------
The above set of calculations determines the system resistance for a specific tota
air volume, Q. Resistances for various air volumes can be computed from the
following relationship:
( v. 2
New system volume \ _ , ,
Calculated system volume] Calculated system resistance
By knowing the computed volume and pressure at the fan, the pressure at other
volumes can be calculated. These points can then be plotted to give a system curve
which describes the system under all flow conditions. The pressure-volume relation-
ship of the fan can be plotted on the same graph. The operating point is the
intersection of the two curves. Figure 9-10 illustrates a typical graph of this type
Note that the above relationship is based on the premise that the pressure losses
through each element of a system vary proportionally to the square of the flow
rate. Certain equipment does not exhibit this relationship, especially some air
pollution control equipment.
Volume, cfm
Figure 9-10. Typical point of operation.
Also, in Example 9-4, air was assumed to be at the standard atmospheric condi-
tions of 70°F and 29.92 in. Hg. Under these conditions, the density of air is 0.075
Ibs/it . Corrections must be made when process conditions vary from these stan-
dards. This is especially true for systems handling hot, saturated air streams
9-24
-------
Fans
In controlling air pollution, it is necessary to move the contaminants from their
point of generation to the control device. Air and gas moving devices using rotary
motion are normally termed fans. Fans are classified as axial or centrifugal
depending on the direction of air flow through the impeller.
In axial flow fans, air moves directly forward through the axis of rotation of the
fan blades (Figure 9-11). A home window or room fan is an example of an axial
flow fan. For industrial use, axial fans are best suited for moving large volumes of
clean air with low static pressures.
5 Air flow
Tube-axial fan
Guide vane
Air flow
Vane-axial fan
Figure 9-11. Axial fans.
9-25
-------
In centrifugal fans, the air is introduced into the center of a revolving wheel or
rotor and exits at right angles to the rotation of the blades (Figure 9-12) Cen-
trifugal fans are more widely used than axial fans, since they can handle dirtier air
streams with higher system resistances.
Scroll side
Outlet
Rim
(shroud, wheel ring
retaining ring,
wheel rim)
Scroll
(housing, volute)
Impeller
(wheel, rotor)
Forward curved
Backward curved
Radial
Airfoil
Figure 9-12. Centrifugal fans.
Two other terms used to describe air movers are blowers or compressors.
Distinguishing between air movers and fans is somewhat arbitrary and is based on
the static pressure they are capable of producing. The term "fan" is generally used
to describe air movers with pressure rises up to about 2 psig (Pollak, 1973). The
term "blower" refers to a centrifugal fan or positive displacement compressor
operating with static pressures in the 2 to 15 psig range. For pressures higher than
15 psig, the term "compressor" is used. Unlike fans, some blowers and compressors
operate by piston movement (positive displacement) rather than rotation.
9-26
-------
Fans
In controlling air pollution, it is necessary to move the contaminants from their
point of generation to the control device. Air and gas moving devices using rotary
motion are normally termed fans. Fans are classified as axial or centrifugal
depending on the direction of air flow through the impeller.
In axial flow fans, air moves directly forward through the axis of rotation of the
fan blades (Figure 9-11). A home window or room fan is an example of an axial
flow fan. For industrial use, axial fans are best suited for moving large volumes of
clean air with low static pressures.
: Air flow
Tube-axial fan
Guide vane
Air flow
Vane-axial fan
Figure 9-11. Axial fans.
9-25
-------
In centrifugal fans, the air is introduced into the center of a revolving wheel or
rotor and exits at right angles to the rotation of the blades (Figure 9-12). Cen-
trifugal fans are more widely used than axial fans, since they can handle dirtier air
streams with higher system resistances.
Scroll side
Outlet
Rim
(shroud, wheel ring,
retaining ring,
wheel rim)
Scroll
(housing, volute)
Impeller
(wheel, rotor)
Forward curved
Backward curved
Radial
Airfoil
Figure 9-12. Centrifugal fans.
Two other terms used to describe air movers are blowers or compressors.
Distinguishing between air movers and fans is somewhat arbitrary and is based on
the static pressure they are capable of producing. The term "fan" is generally used
to describe air movers with pressure rises up to about 2 psig (Pollak, 1973). The
term "blower" refers to a centrifugal fan or positive displacement compressor-
operating with static pressures in the 2 to 15 psig range. For pressures higher than
15 psig, the term "compressor" is used. Unlike fans, some blowers and compressors
operate by piston movement (positive displacement) rather than rotation.
9-26
-------
Axial Fans
Axial fans are classified as tube-axial or vane-axial (see Figure 9-11). Axial fans
can be designed to handle a wide range of air volumes, ranging from a few
hundred cfm to 500,000 to 600,000 cfm.
Tube-axial fans are the simplest fan type, consisting of a propeller mounted in a
tube (Figure 9-11) and therefore, are usually referred to as a duct fan. If the pro-
peller blades are mounted outside a tube (roof or room vent) it is referred to as a
propeller fan.
Tube-axial fans operate at static pressures as high as 4 inches of water and give a
maximum mechanical efficiency (output flow power/input brake horsepower) of
75 to 80% (Summerell, 1981).
The vane-axial fan is similar to the tube-axial, except it has been aero-
dynamically designed to operate at higher efficiencies (see Figure 9-11). One
difference is that it contains guide vanes that correct the helical motion of the air
leaving the blade tip. This reduces the amount of turbulence produced and
increases mechanical efficiency to a maximum of 85% (Summerell, 1981). High
efficiency vane-axial fans are more efficient than comparable centrifugal fans
(Crocker, 1981). They have been used for energy conservation in Europe for a
number of years.
Vane-axial fans normally consist of short, stubby blades mounted on a large
hub, which also helps to reduce turbulence. Single stage vane-axial fans can
operate at pressures as high as 6 to 9 inches of water. Double stage fans are
capable of much higher pressures, but are seldom used (Alden, 1970).
Centrifugal Fans
Air entering a centrifugal fan is rotated between the blades and ejected at a 90°
angle to the blade. The centrifugal force compresses the air which gives added
static pressure to the air. Centrifugal fans are enclosed in a scroll-shaped housing
which helps convert kinetic energy to static pressure. Air rotating between the fan
blades is compressed in the fan scroll which increases the static pressure. Cen-
trifugal fans are classified by blade configuration as radial, forward curved,
backward curved or airfoil.
Radial or Straight Blade
Radial or straight blade fans physically resemble a paddle wheel with long radial
blades attached to the rotor and are the simplest design of all centrifugal fans
(Figure 9-12). This enables most radial blade fans to be built with great
mechanical strength and be easily repaired.
These fans can be used in a variety of applications, especially handling dirty or
sticky materials at high pressure. The flat blades tend to resist abrasion better than
other designs. They can handle high static pressures, to 60 inches water gage and
beyond (Summerell, 1981). However, these fans are the least efficient of the
centrifugal type. A typical efficiency is 65%, but an efficiency of 12% is not too
uncommon (Perry, 1976).
9-27
-------
Forward Curved
Forward curved fans have both the heel and the tip of the blade curve forward in
the direction of rotation (Figure 9-12). Blades are smaller and spaced muTc^r
together than in other blade designs. Air leaves the tip of the blades at a greateT
velocity than the wheel-tip speed. Therefore, these fans are capable of JS£L
a large volume of gas for a minimum wheel diameter. ^errin^
Forward curved blade fans are generally not used when dust or sticky materials
drc present. v^ont cunm 3,n ts e3.silv sfiflc tn fli^v^io/-? j •
y tiic oidCios cinci C3u.sc ooer3.tioTi3]
. These fans are used in low pressure (4 inches water gage), high air
entllafinn ar»nli/-o*i/-i»> A/fn..;__ „_ i •. «_ . °
Backward Curved or Backward Inclined
Backward curved or backward inclined fans have blades inclined in opposition to
he direction of rotation. This feature causes air to leave the tip of the Wade at I
lower velocity than the whppl t\n cr»»»/4 TU- • , ; ^«uc
-------
100
8 10 12 14 16 18 20 22 24 26
Volume, cfmx 1000
Figure 9-13. Typical performance curve for a backward
curved centrifugal fan.
Performance curves can be used to select the proper fan for a particular job.
The point of operation (or rating) is determined by the intersection of the static
pressure and the system flow resistance lines (see Figure 9-10). Curves like those of
Figure 9-10 are used to determine if the operating point is in a desirable location.
For example, the system curve should intersect the static pressure line so that
operation is near the maximum efficiency range and is in a low power requirement
range. Also, the system curve should intersect the pressure curve in an area with a
steep slope. This gives good flow control since large pressure changes bring small
variations in air volumes (Figure 9-14). If the pressure curves were flat, a small
pressure change would be accompanied by a large volume change. These flow con-
siderations are important since system pressure usually is somewhat variable and
the operation point is determined from design data which may vary from actual
operation.
9-29
-------
Effect of varying system pressure
Volume
Poor selection
Fan with flat pressure curve gives wide
volume variation with pressure change.
Volume
Good selection
Fan with steep pressure curve gives
small volume variation with pressure.
Effect of fan size
a,
_y
Volume
Poor selection
Small fan used with system curve
crossing fan too far to right of peak.
Excessive horsepower.
Low efficiency.
Volume
Good selection
Large fan used with curve crossing
fan curve to right of peak.
Low horsepower.
High efficiency.
Source: Permission granted by Committee on Industrial Ventilation, American
Conference of Governmental Industrial Hygienists, Box 16153, Lansing, MI
48901. Taken from 14th Edition (1976) of Industrial Ventilation.
Figure 9-14. Use of performance curves for fan selection.
To help in the actual selection of fan size and speed, fan manufacturers have put
together multi-rating tables. These tables give the performance of a particular fan
size for its range of capacities. Table 9-5 is a typical fan rating table. The fan size
and dimensions are usually listed at the top of the table. Values of static pressure
are arranged as columns which contain the fan speed and horsepower required to
produce various volume flows. The point of maximum efficiency at each static
pressure is usually underlined or printed in a special type.
9-30
-------
Table 9-5. Typical fan rating table.
Wheel style: backward—inclined
Wheel diameter: 50Vi in.
Maximum fan speed: 1134 rpm
Performances underlined are those at maximum efficiency
cfm
5727
6873
8018
9164
10309
11455
12600
13746
14891
16037
17182
18328
19473
20619
21764
P*
OV
1250
1500
1750
2000
2250
2500
2750
3000
3250
3500
3750
4000
4250
4500
4750
y>
rpm
216
236
250
271
293
315
337
360
382
405
429
451
474
497
520
in.
BHP
0.74
0.97
1.27
1.63
2.12
2.72
3.46
4.39
5.43
6.66
8.08
9.64
11.46
13.51
15.82
1 in.
rpm
278
291
305
320
338
356
377
399
421
442
463
486
510
532
556
BHP
1.34
1.67
2.05
2.49
3.05
3.65
4.39
5.25
6.25
7.48
8.97
10.61
12.47
14.55
16.82
IV* in.
rpm
330
339
352
366
381
396
413
430
451
473
496
517
539
560
584
BHP
2.01
2.42
2.87
3.42
4.06
4.76
5.58
6.48
7.52
8.67
10.05
11.61
13.47
15.60
17.98
2 in.
rpm
378
384
393
405
419
432
448
465
481
501
521
543
565
587
608
BHP
2.75
3.24
3.76
4.35
5.06
5.88
6.81
7.82
8.93
10.16
11.54
12.99
14.85
16.82
19.17
2VSin.
rpm
421
424
430
441
453
468
482
496
512
529
547
566
587
610
632
BHP
3.50
4.06
4.69
5.36
6r14
7.03
8.04
9.16
10.35
11.69
13.18
14.78
16.56
18.54
20.77
Sin.
nap
460
462
467
475
486
499
514
527
542
557
574
591
610
630
652
BHP
4.28
4.91
5.66
6.40
7.26
8.19
9.31
10.53
11.80
13.25
14.81
16.49
18.35
20.40
22.59
Source: Bayler Blower Co.
Note: performances based on standard air density of 0.075 Ibs/ft8.
In specifying the proper size fan, it is desirable to know the horsepower require-
ment for the system. Two terms are used when describing horsepower: air
horsepower (AHP) and brake horsepower (BHP). Air horsepower is the power out-
put of a fan if the fan were 100% efficient. Since no fan is 100% efficient the
more important term is the brake horsepower. Brake horsepower equals the air
horsepower times the efficiency of the fan. Knowing the efficiency of the fan, brake
horsepower can be calculated from:
(Eq. 9-12)
BHP=
r
6356 x
-------
It IS also important to note that most multi-rating tables and performance curves
are based on air at standard conditions (70°F and 14.7 psia). When air is not at
standard conditions, volume, system pressure, and horsepower must be corrected.
These corrections are proportional to the change in density at the operating
temperature and pressure.
Fan Laws
Fan laws are used to predict how a change in one of the performance variables
(size, speed, air density, volume, system pressure, power, efficiency, and sound
level) affects the others. Fan laws are based on the premise that, if two fans are
geometrically similar (homologous) their performance curves have the same
shape. Therefore, if the two fans operate at the same point of rating, their effi-
ciencies are equal and the ratio of all other variables are interrelated.
The fan laws can be written in as many as ten different ways depending on
which variables change. The following relationships are some of the more useful of
the fan laws.
For a change in fan speed (rpmi — rpm2):
• the flow rate (Q) varies directly with the fan speed ratio
Qi _ rpmi
Qf rpm2
• the power (g>) varies with the cube of the fan speed ratio
rpn
• the static pressure (p.) varies with the square of the fan speed ratio
p.i\ /rprnA*
— = -t —
—
PV \rpm2/
For a change in fan size (wheel diameter dwl— dw2):
• the flow rate varies with the cube of the wheel diameter
0?
• the power varies with the fifth power of the wheel diameter
"A V
• the static pressure varies with the square of the wheel diameter
9-32
-------
For a change in gas density (QI~~QI)
• the flow rate is constant
C*
• the power varies directly with the density ratio
'3$
f_X n Q «
• the static pressure varies directly with the density ratio
P'i 61
T.... ^H ii
P»Z 62
The fan laws can be used to construct performance curves or to determine new
operating conditions as process variables change. It is important to note that these
fan laws are based on the restriction that the fan at the new conditions will operate
at the same point of rating (mechanical efficiency). Therefore, the complete set of
operating conditions should be calculated for any change.
Example 9-5
A fan operating at a speed of 1474 rpm delivers 10,200 cfm at 4 in. static pressure
and requires 8.85 brake horsepower (BHP). What will be the new operating condi-
tions if the fan is speeded up to 2000 rpm.
Solution:
The new flow is:
Q,
rpm
= 13,840 cfm
The new static pressure of the system is:
Pi? = /rPmA 2
. . /2000\2
., = 4 in. -
\1474/
= 7.4 in.
9-33
-------
The horsepower required is:
# \rpm
«-••« -*
Note the drastic increase in horsepower for increasing the fan speed.
Example 9-6
The exhaust system for a rotary dryer is designed to deliver 12,000 cfm of air at
600 °F and 4 in. H2O static pressure. Fan speed is 630 rpm and requires 13
horsepower. If the dryer has been down for repairs, compute the horsepower
required to pull the same amount of air at ambient conditions.
Given: density of air at 70°F= 0.075 lb/ft3
density of air at 600°F= 0.0375 lb/ft3
Solution:
The horsepower is:
Note: this would require a motor twice the size as needed for normal opera-
tion. An alternative would be to install dampers until the system reaches operating
temperature.
Noise
Some fan systems can be objectionably noisy and cause problems in the work area
or at neighboring residences. Fan manufacturers can usually supply data on the
expected sound level generated by a particular fan. However, it is extremely dif-
ficult to determine the maximum noise level of a fan prior to actual operation.
The sound power generated by a fan depends on the flow, fan-pressure level, and
impeller type and configuration (Pollak, 1973). In general, axial fans are noisier
than centrifugal fans.
A number of possible methods can reduce the noise level of a fan.
1. Apply acoustical insulation to the fan housing or ductwork. A housing
enclosure around the fan unit could be considered if the noise level is severe.
2. Use a venturi inlet to smooth the flow of air into the fan. This can lower the
noise level by reducing turbulence in the inlet of the fan.
3. Install absorption type silencers on the inlet and/or outlet of the exhaust.
These silencers are made for insertion in round or rectangular ducts, in stan-
dard or special materials, and with special acoustical fills for corrosive
atmospheres. These devices do add to the pressure drop in the system.
9-34
-------
The above noise reduction techniques are for fans that are designed and
operated properly. The point of minimum noise level of a fan is close to the point
of maximum efficiency. If a fan is not operating properly, the above solutions may
not cure the problem or be the most effective solution. Putting a fan in proper
balance can be the most effective solution.
References
Alden, J. L. and Kane, J. M. 1970. Design of Industrial Exhaust Systems. New
York: Industrial Press Inc.
Crocker, B. B. 1980. Fans in Kirk-Othmer Encyclopedia of Chemical Technology,
Vol. 9. New York: John Wiley and Sons.
Cheremisinoff, P. E. and Cheremisinoff, N. P. 1976. Calculating Air Flow
Requirements for Fume Exhaust Hoods. Plant Engineering. 111-114
(February 19).
Danielson, J. A. ed. 1973. Air Pollution Engineering Manual. Research Triangle
Park, NC. U.S. Environmental Protection Agency.
Industrial Ventilation 14th ed. 1976. Lansing, Mich. American Conference of
Governmental Industrial Hygienists.
Perry, R. E. 1976. The Operation, Maintenance and Repair of Industrial Centri-
fugal Fans. Combustion. 47:7-17 (February).
Pollak, R. 1973. Selecting Fans and Blowers. Chem. Engr. 80:86-100 (January 22).
Summerell, H. M. 1981. Consider Axial Flow Fans When Choosing a Gas Mover.
Chem. Engr. 88:59-62 (June 1).
9-35
-------
Appendix A
Common International System of Units (SI)
Quantity (1)
length
area
volume
speed or velocity
acceleration
rotational frequency
mass (5)
density
force
movement or force (6)
pressure (or vacuum)
stress
viscosity (dynamic)
viscosity (kinematic)
energy, work, or
quantity of heat
power, or heat flow
rate
temperature, or tem-
perature interval
Some common units
kilometer
meter
centimeter
millimeter
micrometer
square kilometer
square hectometer
square meter
square centimeter
square millimeter
cubic meter
cubic decimeter
cubic centimeter
meter per second (12)
kilometer per hour (4)
meter per second squared
revolution per second
revolution per minute (4)
megagram
kilogram
gram
milligram
kilogram per cubic meter
kilonewcon
newton
newton meter
kilopascal
pascal
megapascal
millipascal second (7)
pascal second
square millimeter per
second (8)
joule (9)
kilowatt hour (10)
kilowatt
watt
kelvin
degree Celsius (11)
Symbol Equivalent Symbol
km
m
cm
mm
fj.m
km1
hml hectare (2) ha
m1
cm*
mm'
m' liter (3) L
dm5
cm' milliliter (3) mL
m/s
km/h
m/s'
r/s
r/min
Mg metric ton t
kg
g
mg
kg/m3 gram per liter g/L
kN
N kilogram kg-m/s2
meter
per second
squared
N-m
kPa
Pa Newton per N/m*
meter
squared
MPa
mPa«s
Pa-s
mm'/s
] Newton N»m
meter
kW.h kilowatthour kWh
kW
W
K
°C
NOTES
(1) Any measurable prop-
erty (such as length, area,
temperature) is called a
quantity.
(2) For land or water area
only.
(3) To be used only for
fluids (both gates and
liquids), and for dry ingre-
dients in recipes, or for
volumetric capacities. Do
not use any prefix with liter
except mtllt.
(4) The symbols for
minute, hour, and day are
min, h, and d, respectively.
(5) Commonly called
weight.
(6) Torque or bending
movement.
(7) 1 mPa-s» 1 cP (cen-
tipoise, which is obsolete).
(8) 1 mmVs=l cSt (cen-
tistokes, which is obsolete).
(9) The unit-multiples
kilojoule (kj) and mega-
joule (MJ) are also com-
monly used.
(10) To be abandoned
eventually. 1 kW«h = 3.6
MJ.
(11) The degree mark ° is
always used in °C to avoid
confusion with coulomb
(C), but never with K for
kelvin.
(12) Second is denoted by s
in SI units.
Source: The American National Metric Council, 1978.
A-l
-------
Appendix B
Conversion Factors
Length
1 inch = 2.54 cm
1 m= 3.048 ft
1 ft = 0.305 m
Mass
1 lb = 453.6g
1 kg=2.21b
Pressure
1 atm=101,S25 Pa
= 760mmHg(0°C)
= 14.7 psia
Force
1 N = 1 kg«m/s2
Energy
1 cal = 4.184J
1 J = 9.48X10'4 Btu
1 Btu = 252.2 cal
Kinematic viscosity
1 mz/s=104 stokes
Power
1 W = 1 J/s
1 hp = 33,479 Btu/hr
Area
1 cm2 = 0.155 in2
1 mz=l 0.764 ft2
Volume
1 cm3 = 0.061 ins
1 ms = 35.31 ft3
1 barrel (oil) = 42 gal
1 ft3 = 7.48 gal
1 ft3 = 28.317 liters
Density
1 kg/ms = 0.06241b/ft3
Dynamic viscosity
1 Pa»s= 1 N»m/s= 1000 centipoise
1 cp = 0.000672 Ib/ffsec
Volume flow
1 ms/s=35.3 ftVsec
1 mVmin=35.3 ftVmin
1 scfm=1.7 mVh
1 gpm= 0.227 mVh
Velocity
1 m/s=3.048 ft/sec
1 mi/hr= 0.447 m/s
Geometry
area of circle = Trr2
circumference of circle = 2 TTT
surface area of sphere = 4 irr2
volume of sphere = 4/3 irr3
area of cylinder = 2 Trrh
B-l
-------
Appendix C
Constants and Useful Information
7T=3.14
Gas constants
R = 0.0821 atm liter/g moNK
= 83.14xl06 g.cm3/s2»g mol«K
= 8.314 N»m/gmol»K
= 0.7302 atm-ftVlb mol-°R
= 1.987 g«cal/g mol»K or Btu/lb mol»°R
Acceleration of gravity
g=32.17 ft/sec2 = 980.7 cm/s2 = 9.8 m/s2
Newton's conversion constant
gc = 32.17 (lbmaMXft)/(lb/orc.)(sec2)
1 Ib mol= 359 ft3 of ideal gas at STP (32°F and 14.7 psia)
1 g mol = 22.4 liters of ideal gas at STP (0°C and 760 mm Hg)
C, for water* 1 Btu/lb»°R«l cal/g °C (at 20°C and 1 atm)
Cp for air « 0.26 Btu/lb-°R«0.26 cal/g °C
viscosity of water, jt= 1 cp = 0.01 g/cm»s (at 20°C and 1 atm)
viscosity of air, /ia = 4.1 x 10'7 Ib»sec/ft2 = 2 x 10'5 N-s/m2 (at 20°C and 1 atm)
density of air= 1.29 kg/m3 = 7.49X 10'2 lb/ft3 (at 20°C and 1 atm)
density of water = 1 g/cm3 = 62.4 lb/ft3 (at 4°C and 1 atm)
1 cubic foot of air weighs 34.11 g
conversion from ppm to g/m3 at STP (273.15 K and 1 atm)
/ e \
ppmxMW —2—
g Vgmol/ 1
dSCm 22 414 — x ID'3 m3 293-15 K l x
gmol 103 liters \273.15 K
C-l
-------
Appendix D
Combustion Calculations
for Theoretical and Excess Air
Requirements
Theoretical or Stoichiometric Air
Consider a generalized fuel with a chemical formula dH^CX where the indices x,
y, z, and w represent the relative number of atoms of carbon, hydrogen, sulfur,
and oxygen respectively. Balancing the chemical reaction for the complete oxida-
tion (combustion) of this fuel with oxygen from air gives:
(Eq.D-1)
0
U .£ 1
Where: Q= heat of combustion
v w\
-2- + z-— N2
The above reaction assumes that:
• air consists of 21% by volume of oxygen with the remaining 79% made up
of nitrogen and other inert gases;
• combined oxygen in fuel is available for combustion, thus reducing air
requirements;
• fuel contains no combined nitrogen, so no "fuel NOX" is produced;
• "thermal NO, " via the nitrogen fixation is small, so that it is neglected in
Stoichiometric air calculations;
• sulfur in fuel is oxidized to SO2 with negligible SOS formation.
For example, to combust methane (CH4) Equation D-l reduces to:
(Eq. D-2) CH4 + 2Oz + 7.53Nz-CO2-!-2H2O + 7.5SNs + Q,
(Eq. D-3) Moles or relative volumes: 1 + 2j^7. 5_3 - 1 + 2_+J . 53
total air required total flue gases
For every mole or standard cubic foot of CH4 burned, the reaction requires 9.53
moles or standard cubic feet of air for complete combustion.
D-l
-------
For mixtures of gaseous fuels it is easier to compute the amount of air required
for each of the constituent compounds, e.g., methane, ethane, ethylene etc
directly, using the constants from Table 3-1 (Chapter 3), and then adding them to
get the total. Further, as the analyses of gaseous fuels are usually available on a
volumetric basis, the volume rather than mass of stoichiometric air is of the most
interest. Thus, for a unit volume of gaseous fuel, say 1 standard cubic foot, the
volume of theoretical air, Vat, also in standard cubic feet is:
(Eq. D-4) V., = 2.38(CO + H2) + 9.53CH, + 16.68C2H6 + 14.29C2H4
+ 11.91C«Hf+ . ..+7.15H2S-4.76O2
where the molecular symbols now represent the volume fractions of the indicated
components, and the numerical coefficients are again found in Table 3-1, but this
time from the "mole per mole of combustibles or ftVft3 combustibles" column.
Should the gas mixture contain other combustible substances not already included
in Equation D-4, these can be added similarly. Absence of a substance means that
its volume fraction is zero and that term will drop out of Equation D-4.
Excess Air
Several concentration corrections have been devised based on the combustion
characteristics of fossil fuels. Excess air is defined as that percentage of air added
in excess of that required to just combust a given amount of fuel. Normally, to
achieve efficient fuel combustion, more air is needed than the stoichiometric
amount, i.e., one carbon atom to two oxygen molecules.
Depending on the amount of excess air, different concentrations of CO2 and
oxygen in the stack gas will result, as shown in Figure D-l.
D-2
-------
-o
v
o
u
•C
O, (natural gas)
O, (#2-#6 oils)
Ot (bituminous coals)
COj (bituminous coals)
COj, (#2-#6 oils)
COj (natural gas)
10 20 30 40 50 60 70 80 90 100
% excess air
Figure D-l. Excess air concentrations of CO, and Ot in stack.
Since the concentration of the pollutants produced in the source could be reduced
by adding more excess air, emission calculations are corrected to a given excess air
condition. A value of 50% excess air has been chosen as a reference condition since
at one time many boilers operated at this condition. When this correction is made,
it also accounts for dilution caused by air leaking in at the preheater or other duct
work.
The expression for percent excess air, as given in EPA Method 3, is:
(Eq. D-5)
%EA =
%0!,-0.5%CO
0.264%N2-(%02-0.5%CO)
100
The derivation of this expression is given in the APTI Course 450 Student
Manual, Appendix D (EPA 450/2-79-006).
D-3
-------
Appendix E
Capital Cost Estimations
of Gaseous Control Equipment
This appendix contains generalized cost data for air pollution control systems
described throughout this manual. These data should be used only as an estimate
to determine system costs. In some cases, the cost of the control device may repre-
sent only a very small portion (<20%) of the total installed cost; in other cases it
may represent the total cost. Variations in the total cost can be attributed to a
number of variable factors such as cost of auxiliary equipment, new or retrofitted
installation, local labor costs, engineering overhead, location and accessibility of
plant site, and type of installation (factory or field assembled).
All cost estimating data, except for condenser data, was adapted from an EPA
publication, Capital and Operating Cost of Selected Air Pollution Control Systems
(1978). All equipment costs except condenser data, are based on a reference date
of December 1977 and estimated to be accurate to within ± 20% on a component
basis, except where noted. For further information on the basis of these cost
estimates the above cited document should be consulted.
Thermal and Catalytic Incinerators
Prices for thermal incinerators including refractory linings, are contained in
Figures E-l and E-2. Catalytic incinerator prices are found in Figure E-3.
Residence times for thermal incinerators are based on 0.5 seconds. From Figure
E-1, the price of a thermal incinerator without a heat exchanger for a gas volume
of 30,000 standard cubic feet per minute and 0.5 second residence time is $99,000.
All gas volumes are measured at standard conditions.
The cost curves for thermal incinerators are based on an operating temperature
of 1500°F. The cost of incinerators operating at other temperatures can be deter-
mined by adjusting the inlet gas flow rates to account for the temperature dif-
ference as follows.
flow rate (scfm) at T _ T+ 460
flow rate (scfm) at T6 TA+ 460
Where: T= new temperature, °F
T6= baseline temperature, 1500°F
(flow rate in scfm)
E-l
-------
o
o
o
-------
Absorbers*
Spray Chambers
The cost of a spray chamber is based on the size, volume of chamber, materials of
construction, and water flow rate. Figure E-4 illustrates typical equipment costs for
a spray chamber.
a
a
s.
c/5
90
80
70
60
50
40
30
20
Note:
Based on chamber velocity of 600 fpm.
Length/diameter = 3.
Effective length =4 to 4.5.
Carbon steel construction.
Does not include refractory.
Spray chamber cost includes vessel
and support rings, platform, ladder,
gratings, spray system, and controls.
50 100 150 200
Inlet gas volume, 1000 acfm
250
Figure E-4. Spray chamber costs vs. inlet gas volume.
Packed and Plate Towers
The cost of absorption towers is shown in Figures E-5 through E-8. The cost of
these towers depends on the size, thickness, and materials of construction of the
vessel since these units are basically custom designed for individual processes and
applications. Figure E-5 illustrates the fabricated cost of a carbon steel vessel shell
in dollars per linear foot plus the cost of two semi-elliptical inlets and outlets. The
height of the vessel is determined by multiplying the number of transfer units by
the height of the transfer units plus some additional height for vapor-liquid separa-
tion at the top of the tower and cleanout at the bottom (typically 2 to 3 ft plus
25% of the tower diameter). The shell thickness is determined from Table E-l for
the expected internal operating pressures and temperatures. In many cases a corro-
sion allowance of 1/8 to 1/4 inch is also added to the minimum thickness for car-
bon steel construction. The cost of the fabricated vessel alone, therefore, includes
the cost of the shell plus the cost of the inlet and outlet. The fabricated cost of a
*Note: Cost estimating data for venturi scrubbers is given in Appendix D of the Air Pollution
Training Institute's Course 413 Manual, Control of Paniculate Emissions. EPA 450/2-80-066.
E-3
-------
skirt which is provided for support of the vessel and flange-type nozzles for shell
penetrations must also be added to the vessel cost. Figure E-6 estimates the cost of
skirts which include a base plate, anchors, an 18 in. diameter access opening, two
reinforced pipe openings and a vent hole. The thickness of the skirt should be
approximately equal to the thickness of the shell. Figure E-7 estimates the
fabricated cost of nozzles for typical shell penetrations such as those required for
the inlet and outlet for the gas and liquid, relief valve connections, pumpout alter-
nates, spray inlets, and manways.
10,00017
•3
ns
O
42
O
OJ
£
CO
Two heads I
(inlet and outlet)
Stainless steel
Type
304
316
Cost adjustment
2.0
2.3
4 6 8 10
Shell diameter, ft
Figure E-5. Fabricated cost of carbon steel vessel.
E-4
------- An error occurred while trying to OCR this image.
-------
" c
O 8s)
Al
2-2
la-
.2 i-
' &
i3 U
l.~
I
3800
3200
2600
2000
140°
800
200
For trays only, cost per tray
must be multiplied by
cost adjustment factor below.
1234
Cost adjustment factor
Stainless steel
support plate or distributor
Item
Support plate
Distributor
Cost factor
1.8
1.3
468
Shell diameter, ft.
10 12
Where: a = carbon steel distributor
b = 304 stainless steel tray
c = carbon steel support plate
d = carbon steel tray
Figure E-8. Cost of tray, support plate, or distributor.
Table E-l. Minimum shell thickness at ambient temperature (carbon steel).
Shell diameter*
(ft)
Internal pressure:
Atmospheric
25 in. WG
50 in. WG
100 in. WG
10 psig
100 psig
200 psig
2
(in.)
1/8
1/8
1/8
1/8
1/8
1/8
1/4
4
(in.)
1/8
1/8
1/8
1/8
1/8
1/4
1/2
6
(in.)
1/8
1/8
3/16
3/16
1/4
3/8
3/4
8
(in.)
1/8
1/8
3/16
1/4
3/8
1/2
1
10
(in.)
1/8
3/16
1/4
5/16
7/16
3/4
1 1/4
12
(in.)
1/8
3/16
1/4
5/16
1/2
3/4
1 1/2
*For corrosion allowance add minimum of 1/8 inch. Thicknesses are for ambient temperatures
and internal temperatures up to 600°F. Thickness correction factors for higher temperatures
are: 1.04 for 700°F; 1.14 for 750°F; 1.35 for 800°F.
E-6
-------
The cost of internal tower equipment such as support plates, trays, and
distributors is shown in Figure E-8. The indicated prices apply to both tray towers
and packed towers and represent the installed cost per item. The cost of trays is
based on quantities of 20 or more. For quantities less than 20, the unit tray cost
must be multiplied by the cost adjustment. The cost per cubic foot of internal
packing is shown in Table E-2.
Table E-2. Cost of tower packing.
Type and
material
Pall rings:
Carbon steel
304 stainless steel
Polypropylene
Intalox saddles:
Polypropylene
Porcelain
Raschig rings:
Carbon steel
Porcelain
1 in.
<$/&')
19
54
14
15
12
17
8
1 1/2 in.
($/ftJ)
15
43
—
—
14
7
2 in.
($/fts)
13
35
8
10
9
9
6
3 in.
($/fts)
—
—
5
8
_
3 1/2 in.
($/ftJ)
_
—
—
:
_
—
The summation of these costs represent the fabricator's shop costs. Added to this
must be the fabricator's cost of engineering, administration costs, and profit as
determined from Table E-3.
Table E-3. Additional costs for fabricator's engineering,
purchasing, administration and profit.
Total cost of fabricated vessel
Less than $5,000
$5,000 to $10,000
$10,000 to $20,000
$20,000 to $30,000
$30,000 to $50,000
$50,000 to $80,000
Over $80,000
Cost factor
0.25
0.23
0.20
0.19
0.18
0.17
0.16
As an example, the design calculations for a carbon steel packed tower indicate
that the tower diameter should be approximately 2 ft with a tower height of
approximately 15 ft. The inlet gas flow rate is 520 cfm and the liquid flow rate is
3 gpm. The tower packing is to be 1 in. carbon steel Raschig rings and the absorber
is to be operated at ambient conditions of temperature and pressure.
Under these conditions, a 0.25 in. plate thickness for the shell may be selected
which would provide adequate allowance for corrosion problems. A suitable inlet
gas velocity for this absorber would be approximately 1800 fpm which would
require a gas inlet and outlet of approximately 8 in. diameter. A suitable liquid
E-7
-------
inlet and outlet pipe for 3 gpm would be approximately 1 in. A manway of 18 to
24 in. diameter would normally be provided at the top and bottom for larger
diameter towers, however, for this small diameter tower, a 12 in. hand hole can be
provided at each end of the tower. A drain in the bottom of the tower should also
be provided for draining, flushing and cleanout.
Using Figure E-5 the cost of the basic vessel is $420 for the heads plus $52 per
foot of shell or $1200. The cost of the skirt, as determined from Figure E-6, is
$1140. The installation of nozzles for the various inlet and outlet diameters results
in a cost of $2580 for three liquid nozzles at $120 each, two gas inlet/outlet (at 8
in.) at $330 each, and two hand holes with blind flanges at $780 each (Figure E-7).
The tower internals would consist of a packing support plate and two distributors
spaced at 5 ft intervals plus the internal packing material. The cost of the support
plate and distributors, as determined from Figure E-8, is estimated to be $1220.
The volume of packing required for a 15 ft tower with a 2 ft diameter is approx-
imately 47 ft3 and the estimated cost of the packing, as determined from Table
E-2, is approximately $800. The total fabricated cost of the packed tower is
therefore $6940. The fabricator's selling price including engineering, administra-
tion costs, and profit, but less taxes and freight, is determined from Table E-3 to
be $6940 x 1.23 or $8540.
Adsorber
Prices for carbon adsorbers are presented in Figures E-9 and E-10 as a function of
total pounds of carbon in the unit. The total or gross number of pounds is deter-
mined by the adsorption rate and the regeneration rate of the carbon for the
emission being controlled. A carbon adsorber will normally be a dual system with
one bed on-line adsorbing while the second bed will be off-line regenerating. A
likely estimate of regeneration time for almost all applications would be between
30 minutes to an hour.
o
o
o
•o
a
o
U
80
70
60
50
40
30
20
10
Note:
These prices are for basic systems:
including adsorber, carbon, blower
or fan, controls and steam regenerator.
Special requirements may double the cost.
These prices are for commercial application.
Add 30% for applications involving
stringent industrial standards.
2468
Weight of carbon, 1000 Ibs.
10
Figure E-9. Prices for packaged stationary bed carbon
adsorption units with steam regeneration.
E-8
-------
Figure E-9 represents packaged units for automatic operation in commercial anc
industrial applications. Commercial applications cover dry-cleaning and solvent
metal cleaning. Industrial applications include lithography and petrochemical
processing. Industrial requirements would increase costs 30% over commercial
requirements. Industrial requirements would include heavier materials for high
steam or vacuum pressure designs and more elaborate controls to assure safety
against explosions and prevent hydrocarbon breakthrough. Figure E-10 presents
custom units, mostly for industrial applications where the gas flow rate exceeds
10,000 acfm. Table E-4 is provided to estimate annualized control cost
requirements for steam, cooling water, maintenance, electricity, and carbon
replacement. The carbon weight used to determine the purchase cost can also be
used to estimate the carbon replacement cost.
§ 700
o
*T 600
(A
8
8 500
u
OJ
•£ 400
« 300
o
•S 200
u
6 100
o
w
(A
_'
-
-
-
-
-
-X
— I—
X
,
— r—
•S
,
r-,
^x
,
— r-
!
'
x
— L_
'
/
I
i — r
,
1
/
,
/^
-
-
-
-
-
, •
Note:
•Instrumentation and
controls included.
40 80 120 160
Total weight of carbon, 1000 Ibs.
200
Figure E-10. Prices for custom carbon adsorption units.
Table E-4. Technical assumptions for estimation of direct operating costs.
Item
Steam consumption
Cooling water
Electricity
Maintenance
Carbon replacement
Assumption
4 Ib per Ib pollutant recovered
12 gal per 100 Ibs steam
5 HP per 1000 acfm
5% of equipment purchase cost
Replace original carbon every five
years
E-9
-------
Condensers
Estimated capital costs are for a complete condenser systein including a shell-and-
tube condenser, a storage tank, a pump, and the necessary piping are presented in
Figure E-ll. The shell-and-tube condenser is a fixed-tube type with 1 inch
diameter tubes that are 8 feet long. Costs are based on December 1979 and include
a 30% contingency allowance. Figure E-ll was taken from an EPA publication as
part of a study for controlling emissions from the Organic Chemical Manufacturing
Industry (EPA, 1980).
10,000 F
o
o
o
s|
2 §
'5> «j
------- An error occurred while trying to OCR this image.
------- An error occurred while trying to OCR this image.
-------
Appendix H
Specific Weight of Dry Air in lb/ft3
for °F and °R and Absolute Pressure
of 29.92 in. Hg.
a!
o
a
o
H
4000
3000
2000
1500
1000
900
800
700
600
500
400
300
200
100
\
\
\
\
\
\
^
\
\
\
\
\
\
\
\
-4000
3OUU
-3000
-2000
-1000
-500
-400
-300
-200
-100
-0
--100
--200
--300
a
•3
4-t
a
i~
o
Q.
U
h
0.01 0.015 0.02
0.03 0.04 0.08 0.10 0.15 0.20
Specific weight, lb/ft3
0.30 0.40
H-l
-------
Appendix I
Use of Psychrometric Chart
for Gas Moisture Calculation
Psychrometry is defined as the science dealing with determination of the properties
of gas and vapor mixtures. Psychrometric charts are graphical representations of
the properties of a gas-vapor mixture. By far, the most important gas-vapor mix-
ture is that of air and water. Many forms of psychrometric charts (or humidity
charts) are available for determining the properties of air and water mixtures.
Figure 1-1 is an example of a psychrometric chart. The following example
illustrates the use of this chart in determining the properties of air at different
temperatures and degrees of saturation.
Problem Statement
A coal fired utility boiler uses a combination spray and venturi scrubber to reduce
particulate matter and SO2 gas to acceptable levels. Flue gas at a flow rate of
98,750 acfm leaves the economizer section of the boiler at S05°F. The flue gases
are ducted into the spray section of the scrubber where the gases are cooled and
saturated, and then they flow into the venturi. A wet bulb/dry bulb temperature
measure of 110°F/305°F was taken of the gases entering the scrubber.
Calculate: 1. saturation temperature and humidity of flue gases entering
the spray section
2. saturation humidity
3. evaporation rate
4. saturated volumetric flow
5. percent gas flow reduction as it passes through the scrubber
system
Solution:
1. The humidity is read from Figure 1-1.
The humidity is determined by starting at the dry bulb temperature of S05°F,
following this line up until it crosses the wet bulb line of 110°F and then
reading the humidity as 0.012 lb water/lb dry air.
The saturation temperature is the wet bulb temperature, 110°F. An air stream
can be cooled and saturated no lower than its wet bulb temperature.
1-1
-------
POUNDS WATER K» it. DRY AM
1 *y -*V .1^ \ I i«i"V
t^frSr-^ Xr>«i-»-^a^y'
vbff£x.x>v ^^y-^x^s -
XXXJK X\ X
-S8 BSSS3 8= =§S 8 I g g 8 §
Source: From Chemical Engineers'Handbook by R. Perry and C. Chilton. ©1973
McGraw-Hill. Used with the permission of McGraw-Hill Book Company.
Figure I-l. Psychrometric chart. Properties of air and water-vapor mixtures from 32° to 600° F.
1-2
-------
2. The saturation humidity of the flue gases is determined by following the curved
wet- bulb temperature line until it hits the saturation line and reading off the
humidity at this point, 0.058 Ib water/lb dry air.
3. The evaporation rate of the scrubbing water is the difference between the
incoming and saturation humidity:
0.058 Ib water/lb dry air
-0.012 Ib water/lb dry air
0.046 Ib water/lb dry air
To determine the water make up rate, first compute the Ib of dry air (flue gas)
coming into the scrubber. The ft3 of flue gas can be converted to Ib of moist air
by reading the volumetric flow from Figure 1-1. For a dry bulb/wet bulb
reading of 305°F/110°F this is:
_ ft3 of moist air
Ib dry air
The total Ib of dry air entering the scrubbing system is:
98,750-^- x _ lb ^ air _ = 5000 lb dry air
min 19.75 ft3 moist air min
The water evaporation rate is:
5000 lb diy air x °-046 lb water x 60min = 13,800
min lb dry air hr hr
4. The saturation volumetric flow rate is read from Figure 1-1 where the wet
bulb line (110°F) hits the saturation line. These are the curved, dashed lines;
the saturation volumetric flow rate is:
. c 0 ft3 of moist air
lo.o -
lb of dry air
5. To calculate the gas flow reduction first calculate the flow leaving the scrubber
system. The volume of gas leaving the scrubber is saturated, therefore:
5000* lb ^ air x 15-8ft3°fm°istair= 79.000 acfm
min lb dry air
w , . 98,750-79,000 on .M
% reduction = — ' - : - = 20.0%
98,750
*Note: The volumetric flow of dry air through the system is constant.
1-3
-------
TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
i. REPORT NO. 2.
EPA 450/2-81-005
4.
7.
9.
12
15
TITLE AND SUBTITLE
APTI Course 415
Control of Gaseous Emissions
Student Manual
AUTHOR(S)
Gerald T. Joseph, David S. Beachler
... , . - « , «.,
PERFORMING ORGANIZATION NAME AND ADDRESS ••• »' ' "'*
Northrop Services Inc.
P.O. Box 12313
Research Triangle Park, NC 27709
. SPONSORING AGENCY NAME AND ADDRESS
U.S. Environmental Protection Agency
Manpower and Technical Information Branch
Research Triangle Park, NC 27711
3. RECIPIENT'S ACCESSIOWNO.
S. REPORT DATE
December 1981
6. PERFORMING ORGANIZATION
8. PERFORMING ORGANIZATION
CODE
REPORT
10. PROGRAM ELEMENT NO.
B18A2C
11. CONTRACT/GRANT NO.
68-02-2374
13. TYPE OF REPORT AND PERIOD COVER
Student Manual
14. SPONSORING AGENCY CODE
SUPPLEMENTARY NOTES ' " '
EPA Project Officer for this Student Manual is R.E. Townsend9 EPA-ERC, MD-20,
Research Triangle Park, NC 27711
The Student Manual is to be used in conducting APTI Course 415 "Control
of Gaseous Emissions". This manual supplements the course lecture material,
presenting detailed discussion on gaseous emission control equipment. The
major topics include: Basic Gas Properties, Combustion, Absorption,
Adsorption, Condensation, and Control of NOX and SO? Emissions. This manual
will assist the reader in evaluating plans for gaseous emission control
systems and in conducting plan reviews.
This manual is intended for use in conjunction wich the Instructor's
Guide (EPA 450/2-81-004) and the Student Workbook (EPA 450/2-81-006) for
APTI Course 415.
U.£ Cr:\'!r! H \7A ? ? 1 6 1
b.lDENTIFIERS/OPEN ENDED TERMS
Student Manual
19. SECURITY CLASS (This Report)
Unclassified
20. SECURITY CLASS (This page)
Unclassified
c. COS AT I Field/Group
13B
51
68A
21. NO. OF PAGES
304
22. PRICE
EPA Form 2220-1 (9-73)
1-4
------- |