&EPA
United States
Environmental Protection
Agency
Industrial Environmental Research EPA-600/7-79-020
Laboratory January 1979
Research Triangle Park NC 27711
Evaluation of Granular
Bed Filters for
High-temperature/
High-pressure
Particulate Control
Interagency
Energy/Environment
R&D Program Report
-------
RESEARCH REPORTING SERIES
Research reports of the Office of Research and Development, U.S. Environmental
Protection Agency, have been grouped into nine series. These nine broad cate-
gories were established to facilitate further development and application of en-
vironmental technology. Elimination of traditional grouping was consciously
planned to foster technology transfer and a maximum interface in related fields.
The nine series are:
1. Environmental Health Effects Research
2. Environmental Protection Technology
3. Ecological Research
4. Environmental Monitoring
5. Socioeconomic Environmental Studies
6. Scientific and Technical Assessment Reports (STAR)
7. Interagency Energy-Environment Research and Development
8. "Special" Reports
9. Miscellaneous Reports
This report has been assigned to the INTERAGENCY ENERGY-ENVIRONMENT
RESEARCH AND DEVELOPMENT series. Reports in this series result from the
effort funded under the 17-agency Federal Energy/Environment Research and
Development Program. These studies relate to EPA's mission to protect the public
health and welfare from adverse effects of pollutants associated with energy sys-
tems. The goal of the Program is to assure the rapid development of domestic
energy supplies in an environmentally-compatible manner by providing the nec-
essary environmental data and control technology. Investigations include analy-
ses of the transport of energy-related pollutants and their health and ecological
effects; assessments of, and development of, control technologies for energy
systems; and integrated assessments of a wide range of energy-related environ-
mental issues.
EPA REVIEW NOTICE
This report has been reviewed by the participating Federal Agencies, and approved
for publication. Approval does not signify that the contents necessarily reflect
the views and policies of the Government, nor does mention of trade names or
commercial products constitute endorsement or recommendation for use.
This document is available to the public through the National Technical Informa-
tion Service, Springfield, Virginia 22161.
-------
EPA-600/7-79-020
January 1979
Evaluation of Granular Bed
Filters for High-temperature/
High-pressure Paniculate
Control
by
Shui-Chow Yung, Ronald Patterson,
Richard Parker, and Seymour Calvert
Air Pollution Technology, Inc.
4901 Morena Boulevard, Suite 402
San Diego, California 92117
Contract No. §8-02-2183
Program Element No. 1NE624
EPA Project Officer: Dennis C. Drehmel
Industrial Environmental Research Laboratory
Office of Energy, Minerals, and Industry
Research Triangle Park, NC 27711
Prepared for
U.S. ENVIRONMENTAL PROTECTION AGENCY
Office of Research and Development
Washington, DC 20460
-------
ABSTRACT
The status and potential of granular bed filter (GBF) tech-
nology for fine particulate control has been critically reviewed
and evaluated with emphasis on high temperature and pressure (HTP)
applications.
Available theroretical models and experimental data have been
evaluated and found to be inadequate for predicting the performance
of industrial GBF systems. Additional experimental data were ob-
tained with a bench scale GBF. These data were used as the basis
for a clean bed performance model based on inertial impaction as
the primary collection mechanism. Predictions were in good agree-
ment with data from industrial GBF systems.
The performance and economics of fixed, continuously moving,
and intermittently moving GBF systems have been evaluated for HTP
applications. At the present stage of development, GBF performance
is neither efficient enough nor sufficiently reliable to satisfy
HTP particulate control requirements. Therefore, cost estimates
are highly speculative and serve mainly to indicate areas where
further development work might substantially reduce capital and
operating costs.
This study shows that GBF systems have potential for HTP
applications but further development and design improvements
will be necessary before these systems can be considered adequate
HTP particulate control technologies. Recommendations are made
for improving the efficiency and reliability of GBF systems.
111
-------
CONTENTS
Page
Abstract
Figures vi
Tables xii
Abbreviations and Symbols xiv
Acknowledgement xviii
Sections
1. Summary and Conclusions 1
Current Technology Evaluation 1
Engineering Design Equations 13
Potential for HTP Applications 16
Conclusions 17
2. Introduction 20
3. Literature Review 22
Patents 22
GBF Studies 27
Theory 63
4. Comparison Between Model Prediction and Data 105
Data Reported by McCain 105
Data Reported by Hood 110
Data Reported by Knettig and Beeckmans Ill
Data by Paretsky, et al 116
Data by Gebhart, et al 116
Conclusions 118
5. Experiment 119
Apparatus 119
Data Analysis 122
Pressure Drop Data 131
IV
-------
CONTENTS (continued)
6. Design Model 135
Mathematical Modeling 135
Comparing Model Predictions with Performance Data. . . 141
7. Present Technology Evaluation 145
Granular Bed Filter Systems 145
Industrial Users and Performances 161
Evaluation 175
8. Potential for HTP Applications 185
Cleanup Requirements . 185
Predicted GBF Performance 188
Gaseous Pollutants 192
Preliminary Cost Estimate 193
Permissible Costs 20°
Summary and Conclusions 203
9. Recommendations fur Future Research and Development . . .205
Efficiency Improvement 205
Bed Cleaning Methods 208
Other Research Recommendations 209
Recommended Research Program 209
References
212
-------
FIGURES
Number Page
1 Filtration of aerosols through bed packed with lead
shot (Thomas and Yoder data) 30
2 Time vs. efficiency - 0.033 m deep granular bed
filtering fly ash (Taub's data) 33
3 Penetration of 1.1 ym aerosol as a function of
superficial gas velocity: 10-14 mesh sand (Paretsky,
et al. data) 36
4 Penetration of 1.1 \im aerosol as a function of
superficial gas velocity: 20-30 mesh sand (Paretsky,
et al. data) 36
5 Experimental results obtained by Gebhart, et al.:
Penetration vs. particle diameter for different flow
velocities and granule size 37
6 Experimenta results obtained by Gebhart, et al:
Penetration vs. particle diameter for different flow
velocities and granule size 37
7 Experimental results obtained by Gebhart, et al. :
Penetration vs. particle diameter for different flow
velocities and granule size 38
8 Experimental results obtained by Gebhart, et al:
Penetration vs. particle diameter for different flow
velocities and granule size 38
9 Gravity effect in granular beds measured with an
upward and downward directed aerosol stream (Gebhart
et al. data) ' 40
10 Gravity effect in granular beds measured with an
upward and downward directed aerosol stream (Gebhart
et al. data) ' 40
11 Capture efficiency (transfer units) vs. bed height
for the grid supported fixed bed (Knettig and Beeck-
mans' data) 41
VI
-------
FIGURES (continued)
Number Paj
12 Capture efficiency (transfer units) vs. bed height
for the screen supported fixed bed (Knettig and
Beeckmans1 data) ...................
13 Effect of mean gravel diameter and particulate load
on granular bed collection efficiency (Miyamoto and
Bonn's data) ..................... 43
14 Effect of gravel layer thickness and particulate
load on granular bed collection efficiency (Miyamoto
and Bonn's data) ................... 43
15 Effect of superficial gas velocity and particulate
load on granular bed collection efficiency (Miyamoto
and Bonn's data) ................... 43
16 Penetration tests on 420 to 710 micrometers plastic
beads by monodispersed polystyrene latex aerosol
(downflow) (Figueroa and Licht's data) ........ 46
17 Effect of bed height and bed granule (bead) size on
the downflow penetration of 0.5 micrometers poly-
styrene latex aerosol particles on plastic beads
(Figueroa and Licht's data) .............
18 Comparison of downflow penetration of polystyrene
latex aerosol particle on 420 to 710 micrometers
plastic beads and -25+40 mesh sand granules
(Figueroa and Licht's data) ............. 4
19 Effect of bed granule (bead) size and flow direction
on the penetration of 0.5 micrometer polystyrene
latex aerosol particles on plastic beads ....... 47
20 Schematic of granular bed filter (Westinghouse setup) 4^
21 Schematic of test equipment (Westinghouse) ...... ^0
22 Electrified packed bed ................ 61
23 Performance characteristics of electrostatically
augmented packed bed (Research-Cottrell data) ....
24 Dendrite initiation, growth and idealization of the
dendrite configuration ......... .......
70
25 Particle deposition on single collector .......
26 Typical isotherms for Langmuir and BET adsorption R_
patterns ............... ........
vii
-------
FIGURES (continued)
Number Pase
27 Schematic representation of fly ash emission
mechanisms ...................... y5
28 Fraction of total fly ash emitted by various
mechanisms as a function of deposit thickness . . . . 95
29 A comprehensive plot of pressure drop in fixed
beds ........................ 97
30 Effect of mean granule diameter on pressure drop
(Miyamoto and Bonn's data) ....... . ...... yy
31 Effect of bed depth on pressure drop (Miyamoto and
Bonn's data) ..................... yy
32 Effect of superficial gas velocity on pressure drop
(Miyamoto and Bohn's data) ..............
33 Comparison of McCain's gravel bed particle collection
data with design equation predictions ........
34 Comparison of Hood's data with predictions by avail-
able design equations ................
35 Experimental and predicted performance of CPC GBF
(A.P.T. data) .................... 112
36 Experimental and predicted performance of CPC GBF
(A.P.T. data) .................... 113
37 Experimental and predicted performance of CPC GBF
(A.P.T. data) .................... 113
38 Comparison of Paretsky, et al. data with predictions
by available design equation ............. 117
39 Comparison of Gebhart, et al. data and predictions
by Paretsky, et al. equation. Collection is in the
diffusion regime ............
40 Schematic diagram of the experimental apparatus . . . 12°
41 Particle size distribution for iron shot ....... 121
42 Experimental particle penetration of a clean granular
bed filter .................. . ... 123
43 Experimental particle penetration of a clean granular
bed filter ...................... 123
viii
-------
FIGURES (continued)
Number Page
44 Experimental particle penetration of a clean
granular bed filter ........ . ........ 124
45 Experimental particle penetration of a clean
granular bed filter ................. 124
46 Experimental particle penetration of a clean
granular bed filter ................. 125
47 Experimental particle penetration of a clean
granular bed filter ................. 125
48 Experimental particle penetration of a clean
granular bed filter ................. 126
49 Experimental particle penetration of a clean
granular bed filter ................. 126
50 Experimental particle penetration of a clean
granular bed filter ................. 127
51 Experimental particle penetration of a clean
granular bed filter ................. 127
52 Experimental particle penetration of a clean
granular bed filter ........... ...... 128
53 Experimental particle penetration of a clean
granular bed filter ................. 128
54 Experimental and predicted pressure drops across
a clean granular bed filter ............. 132
55 Experimental and predicted pressure drops across
a clean granular bed filter ............. 132
56 Experimental and predicted pressure drops across
a clean granular bed filter .............
57 Experimental and predicted pressure drops across
a clean granular bed filter ............. 133
58 Experimental and predicted pressure drops across
a clean granular bed filter ............. 134
59 Diagram of granular bed showing impaction concept. . 136
60 Kp vs . n for round jet model for particle collection
in a GBF ...................... 139
ix
-------
FIGURES (continued)
Number Page
61 Kp vs. n for round jet 140
62 Efficiency vs. inertial impaction parameter for
comparison 140
63 Comparison of McCain's gravel bed particle collection
data with design equation predictions 142
64 Comparison of Hood's data with predictions by avail-
able design equations 142
65 Experimental and predicted performance of CPC GBF
(A.P.T. data) 143
66 Experimental and predicted performance of CPC GBF
(A.P.T. data) 143
67 Experimental and predicted performance of CPC GBF
(A.P.T. data) 144
68 Dorfan impinge filter 147
69 Consolidation Coal Company filter 148
70 Combustion Power Company "dry scrubber" 151
71 The integral cyclone model of the "dry scrubber". . . 152
72 Possible design for Squires panel bed filter 153
73 Lurgi MB filter 155
74 Rexnord gravel bed filter 157
75 Ducon granular bed filter 159
76 Filter element 160
77 Collection cycle 162
78 Cleaning cycle 162
79 Experimental grade efficiency curve of a Rexnord
gravel bed filter (McCain, 1976) 165
80 Experimental penetration curves for CPC dry scrubber
(Hood, 1976) 168
-------
FIGURES (continued)
Number Page
81 Continuous moving bed GBF ............... 169
82 Experimental grade penetration (CPC data for nominal
configuration) .................... 172
83 Experimental grade penetration (CPC data for thick
bed configurations) .................. 172
84 Experimental grade penetration (CPC data for thick
bed configuration) .................. 173
85 Experimental grade penetration (CPC data for short bed
configuration) .................... 173
86 Experimental grade penetration (CPC data for small
granule configuration) ................ 174
87 Fractional penetration curve for Ducon granular bed. .176
88 Fractional penetration for Ducon GBF (from Bertrand,
et al., 1977) ..................... 176
89 Predicted GBF performance ............... 189
90 Particle size distributions from Exxon miniplant
fluidized bed coal combustors
91 Fixed bed GBF ..................... 195
92 Summary G.E. comparison cycles (cost of electricity
at an assumed capacity factor of 65%)
202
XI
-------
TABLES
Number Page
1 Summary of experimental investigations of GBFs. ... 2
2 GBF evaluation summary ................ 9
3 Available equations for the prediction of particle
collection in a granular bed ............. 14
OQ
4 Results of Hanford sand filter tests ......... "°
32
5 Summary of Zahradnik, et al. data ..........
6 Slopes and intercepts of fitted lines ........
7 Experimentally determined properties of the granular 45
materials (Figueroa, 1974) ..............
8 Westinghouse GBF data, series #2 ...........
9 Westinghouse GBF data, series #3 ........ . . . 52
10 Westinghouse GBF data, series #5 ........... 53
11 Westinghouse GBF data, series #6 ........... 54
12 Westinghouse GBF data, series #7 ........... 55
13 Data obtained from Strauss and Thring granular bed
filter study ..................... 57
14 Available equations for the prediction of particle
collection in a granular bed ............. 107
15 Comparison of Knettig and Beeckmans' data and
predictions ..................... 115
16 Pressure drop and collection efficiency for 1.09 ym
diameter particles at a superficial gas velocity of
50 cm/s ......... . .... 7 ..... . . .129
17 Pressure drop for 50% collection of 1.1 ym diameter
particles ...................... 130
18 Measured and calculated void fraction of the granular
bed ......................... 131
xii
-------
TABLES (continued)
Number Page
"^^"^™"^^^"*^^^™ _^^^«^b«v
19 Typical operating characteristics of Consolidation
Coal Company granular bed filters ........... 149
20 Summary of Rexnord gravel bed users
21 Design specifications of the system as tested ..... 166
22 Test parameters for CPC moving bed filter (from Wade,
et al., 1978) ..................... 171
23 CPC moving bed filter overall penetration correlation. 171
24 Granular bed filter performance (from Bertrand, et al.177
1977) .........................
25 Conditions for high temperature and pressure
particulate collection ................ 186
26 Summary of EGAS phase II performance and cost results.
Xlll
-------
ABBREVIATIONS AND SYMBOLS
a = effective mass transfer area, cm2/cm3
a = a constant, cm2/dyne
A = fraction of the projected area of a single collector
p particle which is available for capturing the aerosol
by settling, fraction
b = stoichiometry constant, dimensionless
B = a constant expressive of the energy of interaction with
the surface, fc/mol
c = concentration of adsorbate in gas phase, g/cm3
c = equilibrium concentration of sorbate in bulk gas phase,
6 g/cm3
c = saturation concentration of the adsorbate, g/cm3
s
C = empirical constant defined by equation (5), dimensionless
C . = inlet particle concentration, g/cm3
C' = Cunningham slip correction factor, dimensionless
&c = granule diameter, cm
d. - jet diameter, cm
d = initial capillary diameter, cm
d = particle diameter, cm or urn
D = effective diffusion coefficient of gaseous reactant in
the residue layer, cm2/s
Dg = gas phase diffusivity, cm2/s
D = particle diffusivity, cm2/s
D - effective particle diffusion coefficient in granule
p layer, cm2/s
Dpore = pore diffusivity, cm2/s
f = fraction factor, dimensionless
f1 = ratio of collector diameter to initial capillary diameter,
dimensionless
F = compression stress, dyne/cm2
g = gravitational acceleration, cm/s2
j = ratio of channel width to packing diameter, dimensionless
xiv
-------
ABBREVIATIONS AND SYMBOLS (cont.)
k = Boltzmann's constant = 1.38 x 10" erg/°K
k = chemical reaction rate constant, fc/mol-s
kg = mass transfer coefficient, cm/s
k re = pore diffusion coefficient, cm2/s
K = equilibrium constant, £/mol
K = proportionality factor, dimensionless
K^ = D'Arcy permeability, cm2
K = inertial parameter, dimensionless
m, = expected particle number in the k'th layer of a dendrite,
number
M = molar weight, g/mol
Mm = granule recirculation rate, kg/kg of gas
N = number of impaction stages, dimensionless
Pt = overall penetration, fraction or percent
Pt, = penetration for particles with diameter "d ", fraction
or percent "
q = amount of sorbate adsorbed per unit weight of sorbent, g/cm3
q = amount of sorbate adsorbed per unit weight of adsorbent in
equilibrium with concentration "C", g/cm3
Q° = number of moles of adsorbate adsorbed per unit weight of
adsorbent in forming a complete monolayer on the surface
r = average pore radius, cm
r = radius of unreacted core, cm
rc = collector radius, cm
rH = hydraulic radius, cm
R = original radius of the reacting particle, cm
R = universal gas constant, J/g mol-s
t = time, seconds
T = absolute temperature, °K
UG = superficial gas velocity, cm/s
urh = actual gas velocity in bed, cm/s
UGI = interstitial gas velocity, cm/s
u- = gas velocity in the jet, cm/s
ut = terminal settling velocity of the particle, cm/s
w
particulate load, g/cm2
xv
-------
ABBREVIATIONS AND SYMBOLS (cont.)
xfi • fraction of reactant "B" converted into product, fraction
X = dust deposit thickness, cm
yA = mole fraction of reactant "A" in gas phase, fraction
y .. = critical trajectory of the aerosol, cm
Z = bed depth, cm
Greek
a = rate of particles approaching a clean fiber per unit
length, cm2/s
AP = pressure drop, dyne/cm2
PG = gas density, g/cm3
p = particle density, g/cm3
p = density of solid, g/cm3
n = overall single granule collection efficiency or single
stage collection efficiency, fraction
nD = single granule collection efficiency by diffusion mechanism,
fraction
TIDI = single granule collection efficiency by direct interception
mechanism, fraction
nrc = single granule collection efficiency by gravity settling,
fraction
UG = gas viscosity, poise
e = bed porosity, fraction
e- = initial void fraction, fraction
$ = relative force field, dimensionless
X = tortuosity factor, dimensionless
0 = angular cylindrical coordinate, measured counterclockwise
from the downstream stagnation point, radian
0 = pressure drop function defined in equation (93)
0£k = rate of increase of "m, "by deposition on particles already
' occupying the k'th layer due to the radial flow component,
cm'1
(r)
k-1 k = rate of increase of"mk"by deposition on particles occupying
* the (k-l)'st layer dui to radial flow component,s~'
f (-1 "^
0k k = rate °^ increase of'W"by deposition on particles already
' occupying the k'th layer, due to the angular flow component,
o
xv i
0
-------
ABBREVIATIONS AND SYMBOLS (cont.)
_T V = rate of increase of"mj"by deposition on particles occupying
* the (k-l)'st layer due to the angular flow component,s~'
T = time for complete reaction, s
£ = internal void fraction of the solid, fraction
Dimensionless Numbers
NN = Nusselt number, dimensionless
Np = Peclet number, dimensionless
ND = Reynolds number, dimensionless
K€
NC = Schmidt number, dimensionless
N, = Sherwood number, dimensionless
xvn
-------
ACKNOWLEDGEMENT
A.P.T., Inc. wishes to express its appreciation for
excellent technical coordination and for very helpful assist-
ance in support of our technical effort to Dr. Dennis C. Drehmel
of the U.S. Environmental Protection Agency.
XVlll
-------
SECTION 1
SUMMARY AND CONCLUSIONS
A number of advanced energy conversion processes under
development require high temperature and/or high pressure removal
of particulate matter to protect critical system components.
These processes must also meet existing and anticipated particu-
late emission standards. Granular bed filter (GBF) systems
have been proposed as suitable devices for removing fine particles
from high temperature and high pressure (HTP) gas streams. How-
ever, theoretical and experimental performance data for GBFs are
sparsely scattered throughout the literature. A thorough survey
and evaluation of current GBF technology has been performed. The
results are presented in this report.
The objectives of this program have been to:
1. Assess current GBF technology for control of airborne
particulate pollutants,
2. Evaluate existing GBF systems,
3. Develop engineering models and design equations to pre-
dict filter performance,
4. Survey present usage problems, and
5. Evaluate the potential of GBFs for high temperature and
pressure applications.
CURRENT TECHNOLOGY EVALUATION
Performance Characteristics
The literature was searched to assess the current
state of GBF technology. Table 1 summarizes the experimental
investigations reported in the literature. Most of the studies
were laboratory scale experiments. The results of these studies
show that the collection efficiency of a GBF is a function of
particle diameter, face velocity, bed depth, granule diameter,
-------
TABLE 1. SUMMARY OF EXPERIMENTAL INVESTIGATIONS OF GBFs.
Investigator
Granular Bed Configuration
Test Conditions
Parameter Studied
Bed Performance
Comments
I. Laboratory Studies
A.
Thomas and
Yoder
(1956)
(a) Bed type: Fixed bed
(b) Bed material: Sand
(c) Average sand grain diameters:
0.161 cm, 0.071 cm,
0.036 cm
(d) Bed height: 3.6 and 7.6 cm
Aerosol: OOP
Aerosol particle size:
0.1 - 1.0 um
Face velocity: 0.11 -
2.2 cm/s
Sand particle size,
particle shape,
face velocity
(a) Penetration varied
with sand particle
size - decreasing
diameter of sand.
(b) Rough and irregular
sand showed higher
collection efficiency
than smooth sand.
(c) .Penetration
varied with sand par-
ticle size, face velo-
city and aerosol parti-
cle size. Efficiencies
ranged from 40 to 99.8%.
Experimental result
demonstrated the
existance of an
aerosol size of
maximum penetration
of about 0.3 vim.
Particle size of
maximum penetration
decreased with in-
creasing face
velocity.
B. McFee and
Sedlet
(1968)
(a) Bed type: Fixed bed
(b) Bed material: Sand
(c) Sand grain diameter:
0.036 to 0.071 cm
(d) Bed height: 15.2 to
76.2 cm
Aerosol: Pu-U-Mo
Alloy fume
Aerosol particle size:
Geometric mean =
0.07 \m, standard
deviation » 2.7.
(Discrete sizes ranged
from 0.02 to 4 pm)
Face velocity: 0.5-68
cm/s
Aerosol grain loading:
approximately 0.11
g/m3
Bed height, face
velocity
(a) Penetration through 15.2
cm of sand varied from
0.08 to 0.57% over
range of face veloci-
ties from 0.5 to f>'.5
cm/s. Maximum penetra-
tion of 0.57% occurred
in range of 20 to 40
cm/s.
(b) Penetration through 76.2
cm of sand varied from
0.004% at 14.2 cm/s to
0.019% at 25 cm/s.
(c) Penetration decreased
with increase in bed
depth up to depth of
30.5 to 45.7 cm, but
relatively small im-
provement occurred for
beds of greater depth.
Experimental results
indicate the per-
centage of penetra-
tion varies with
aerosol particle
size, face velocity,
bed depth and degree
of packing of bed.
Experimental obser-
vations confirm and
extend finding of
Thomas and Yoder.
continued
-------
TABLE 1. (CONTINUED)
Investigator Granular Bed Configuration Test Conditions
Parameter Studied Bed Performance
Comments
(d) Maximum penetration
occurs at lower face
velocity with in-
creasing bed depth.
(e) Pressure drop through
sand varied with bed
dspth and face velo-
city and ranged from
1.4 to 259 cm W.C.
C. Ducon Bed type: Fixed bed with re- Aerosol:
Company verse gas flow cleaning. a) Iron oxide from
(Avco.Inc.) oxygen-lanced
(1969) electric arc
furnace
b) Iron oxide from
oxygen- lanced
open hearth
furance
c) Nickel ore
d) Fly ash
e) Talc dust
f) Plastic dust
Face velocity:
26 cm/s
D. Kovach and (a) Bed type: Fixed bed Aerosol: OOP, fly
Hannan (b) Bed material : Carbon ash
(1970) (c) Granule diameter: Aerosol particle
Aerosol type, inlet (a) Collection efficiency
grain loading ranged from 98 to
(4.6 to 11.4 99.9%.
g/ra3) (b) Pressure drop: 10.1 to
IS. 2 cm W.C.
Aerosol type, (a) Penetration varied
face velocity, with bed granule
granule size size-decreasing
Collection efficiency was
found to be higher for
the finer sized iron
oxide aerosols than for
coarse fly ash. Indi-
cates that physical
characteristics of aero-
sol are an important
factor in performance of
filter-possibly with
regard to agglomeration
behavior.
Penetration of fly ash
aerosol was less than
that for OOP aerosol .
0.4 to 0.16 cm
(d) Bed height: 38 cm
size: fly ash -
5 to 100 um
Face velocity:
5 to 102 cm/s
with decreasing
granule size
(b) Penetration varied
with face velocity
However, particle size
of fly ash varied over
a wide range, and effect
may be due to both
changes in particle
size and aerosol type.
continued
-------
TABLE 1. (CONTINUED)
Investigator
Granular Bed Configuration
Test Conditions
Parameters Studied
Bed Performance
Comments
E. Zahradnik, (a) Bed type: Shaft filter (bed
et al. of slowly falling
(1970) granules through which
gas stream passes
horizontally)
(b) Bed material: Alkalized
alumina (1/16 in.
beads)
(c) Bed width: 30.5 cm
Aerosol: Fly ash
Aerosol particle
size: 11% 13.7 urn
Gas velocity:
Up to 28 cra/s
Fly ash loading:
4.6 to 6.9 g/ra3
Gas mass velocity
(a) Collection efficiency:
99% at flow rate of
0.88 m'/rain and gas
temperature of 400 F.
(b) Pressure drop: S.I cm
W.C.
Primary intent of study
was to investigate
simultaneous removal
of fly ash and S02
by shaft-filter system.
F. Squires
(1970)
(a) Bed type: Fixed bed with
intermittent movement
of solids
(b) Bed material: Sand
(c) Bed width: 3.5 cm
(d) Sand particle diameter:
approximately 0.05 cm
Aerosol: Fly ash
Face velocity: 6 cm/s
Fly ash loading: 4.6
to 9.2 g/m3
Puffback cleaning
(a) Collection efficiency:
> 99%
(b) Pressure drop: 1.3
cm W.C.
High collection effi-
ciency may be due in
part to preagglomeration
of fly ash before enter-
ing filter
G. Paretsky (a) Bed type: Fixed bed
(1972) (b) Bed material: Sand
(c) Sand particle diameter:
10 to 14 and 20 to
30 mesh (approximately
0.15 cm and 0.07 cm)
(d) Bed height: 3.7 to 19.2 cm
Aerosol: Polystyrene
Aerosol particle
size: 1.1 urn
Face velocity: 0.3
to 80 cm/s
Face velocity, bed
height, granule
(a) Penetration at a given
face velocity de-
creased as the gran-
ule size decreased
Results in agreement
with those reported
by other investiga-
tors.
H. Gebhart.- (a) Bed type: Fixed
et al. (b) Bed material: glass beads
(1973) (c) Bead diameter: 0.4, 0.16,
0.05, and 0.0185 cm
(d) Bed height: 10.2, 20.5!
and 41 cm
Aerosol: Polystyrene
latex
Aerosol diameter:
0.1 to 2 pm
Face velocity: 0.29
to 6.6 cm/s
Granule diameter,
aerosol diameter,
face velocity,
bed height, and
direction of gas
flow
practical interest.
(a) Downward gas flow gives Gas velocity is too
lower penetration than low to be of any
upward gas flow.
(b) M¥£asional collection
and gravity settling
decrease with increas-
ing gas velocity
(c) Smaller granules give
higher efficiency
(d) Deeper bed gives
higher efficiency
(e) There exists a maximum
penetration which is
shifted to larger par-
ticles with decreasing
bead diameter
continued
-------
TABLE 1. (CONTINUED)
Investigator
Granular Bed Configuration
Test Conditions
Parameters Studied
Bed Performance
Comments
I. Knettig and
Beeckmans
(1974).
(a)
(b)
(c)
(d)
(e)
Bed type: Fixed and fluidized
Bed material: Glass beads
Bead diameter: 0.043 cm
Bed height: 1-12 cm
Bed support: Screen or grid
Aerosol: Uranine
methyl ene blue
Aerosol diameter
0.8, 1.6 and 2
Face velocity: 8
and 11.2 cm/s
and
.9 um
.2
Aerosol diameter,
support and face
velocity
(a)
(b)
Linear relationship
between collection
efficiency and bed
height
Collection efficiency
per unit volume of
bed increased with
both aerosol diameter
and superficial gas
velocity.
J. Miyamoto and
Bonn (1975)
(a) Bed type: Fixed
(b) Bed material: Gravel
Aerosol: Ammonium
chloride
Aerosol diameter:
0.1-3 um
Particulate load,
on collection
efficiency
(a) Collection efficiency in- Results show the trend
creased with increasing
particulate load.
(b) Thicker bed has higher
initial collection
efficiency but little
effect when there is
filter cake.
(c) Smaller granules have
higher collection
efficiency and have
sharper increase in
collection efficiency
with particulate load.
(d) Higher-face velocity
decreases the rate
of increase in effi-
ciency with particulate
load.
of cake filtration.
continued
-------
TABLE 1. (CONTINUED)
Investigator
Granular Bed Configuration
Test Conditions
Parameters Studied
Bed Performance
Comments
K. Figueroa
(1975)
L. Westinghouse
(Ciliberti,
1977)
(a) Bed type: Fixed and fluidized
(b) Bed material: Plastic beads,
sand
(c) Granule diameter, plastic
beads, 339 and 495 urn
diameter, sand - 702 pm
(a) Bed type: Fixed bed with
reverse gas flow
cleaning
(b) Bed material : Sand
(c) Sand particle size:
-15 +30 mesh
1 and 2 urn diameter
methylene blue, 0.5,
1.1 and 2.0 diameter
polystyrene latex
10 urn limestone
dust
Granule size,
bed height, face
velocity, flow direc-
tion, aerosol size
Face velocity
and bed depth
(a) Due to electrostatic
charges plastic beads
exhibited higher
collection effi-
ciency.
(b) Highest collection
efficiencies were
obtained on the
largest aerosol dia-
meter, with the deepest
bed of the finer
granules.
(a) Filter performance
improves as dust
accumulates in the bed.
Studied low gas
velocity region.
Experiment not well
controlled.
M. Schmidt,
et al. (1978)
(a) Bed type: Fixed bed
(b) Bed material: Polystyrene
beads, quartz gravel
(c) Granule diameter:
polyethylene beads- 33 mm
quartz gravel - 5 mm
Aerosol: DBP and
room aerosol
Face velocity: 15.2,
45.8, and 101 cm/s
Face velocity
granule size,
and aerosol size
(a) At constant face velocity,
the grade efficiency curve
exhibits a minimum at a
particular particle
diameter.
-------
granule shape, aerosol type and packing density or bed porosity.
Other general results are listed below.
1. At a constant face velocity, the grade penetration curve
exhibits a maximum penetration at a particular aerosol diameter.
The particle diameter corresponding to maximum penetration de-
creases with increasing face velocity and with increasing bed
depth.
2. Particle penetration decreases as the granule diameter
is decreased, as packing density is increased, and as bed depth
is increased,
3. Collection efficiency of a GBF composed of rough, irregu-
lar granules is higher than that of beds composed of smooth
granules.
4. Aerosols which tend to agglomerate are more readily col-
lected.
5. Surface and internal filter cakes tend to improve (the.
collection efficiency of the GBF.
6. There is a maximum particle loading that can be retained
in the bed. Above this maximum loading, the collection efficiency
of the GBF declines.
7. The collection efficiency of the GBF can be increased by
the use of augmenting forces. The filtration efficiency of the
GBF can be enhanced by imposing an electric field or a magnetic
field on the filtration medium. A.P.T. is performing laboratory
experiments with charged and uncharged GBFs collecting charged and
uncharged particles. Preliminary results revealed that the collec-
tion efficiency of the GBF improved significantly without any in-
crease in pressure drop when the bed and/or particles were charged,
Present GBF Systems
A granular bed filter is any filtration system utilizing a
non-fluidized bed of discrete granules or particles as the fil-
tration medium. A variety of types of GBFs are identified in the
literature.
GBFs may be classified according to the bed structure or
according to the cleaning method. With respect to the bed
-------
structure, GBFs may be classified as fixed bed, continuously moving
bed, or intermittently moving bed filters. The advantages and
disadvantages of each bed structure are summarized in Table 2.
The following is a more detailed discussion.
Moving Bed Filters -
The continuously moving bed filter is usually arranged in a
cross-flow configuration. The bed is a vertical layer of granular
material held in place by louvered walls. The gas passes hori-
zontally through the granular layer while the granules and collected
dust move continuously downward and are removed from the bottom.
The dust is separated from the granules by mechanical vibration.
The cleaned granules are then returned to the overhead hopper and
the panel by a granule recirculation system.
The Combustion Power Company's dry scrubber is an example
of a continuously moving bed filter. The performance of this
device has been reported by Wade, et al. (1978). They conducted
extensive cold flow tests to investigate the effects of bed depth,
granule diameter and other parameters on the collection efficiency.
In general the CPC moving bed filter was found to be capable
of particulate removal efficiencies in excess of 98% for particles
in the 1 to 10 ymA diameter range. Submicron particles were
collected at an efficiency in excess of 90% in cases with high
velocities, high inlet particle loadings, and low granule rates.
Beds with larger thickness to granule diameter ratios were most
effective in the capture and retention of particles in the 2 to
5 ymA diameter range. Also, intermittent granule movement was
shown to improve efficiency by a few percent.
High temperature tests of the moving bed filter are planned.
No high temperature data are available at this time.
The major advantage of the moving bed filter design is that
the bed granules are removed and cleaned out of the primary gas
stream. This enables efficient cleaning and relatively steady
collection efficiency. Also it is not necessary to isolate fil-
ter units during cleaning so that the total filter area open to
gas flow is available for filtration at any time.
-------
TABLE 2. GBF EVALUATION SUMMARY
GBF Type
FIXED BED
MOVING BED
INTERMITTENT BED
Advantages
1. No granule recirculation.
2. Lower operating cost.
1. External separation
of granule and dust.
1. External separation of
granule and dust.
Disadvantages
1. Plugging of retaining grids
2. Particle seepage through
bed during cleaning cycle.
3. Fluidization redisperse
fine dust during cleaning.
4. Ineffective bed cleaning
causes particle buildup
in bed.
5. HTP valving required for
reverse air cleaning.
1. Erosion of retaining
grids and transport
system.
2. Particle reentrainment
in moving bed.
3. Granule recirculation
may cause high operating
cost.
4. Difficult to form a
filter cake in moving
bed.
5. Needs a granule recir-
culation and granule/
dust separation system.
No suitable mechanical
device identified. Trans
port by pneumatic means
needs large quantity of
compressed air and ener-
gy to heat the transport
air to bed temperature.
1. Low gas capacity can
cause high capital cost.
2. Erosion of retaining grids
and transport systems.
3. Granule recirculation may
cause high operating cost.
4. Surface cake must be found
to avoid bed plugging
problem.
5. HTP three-way valving
required for cleaning air.
6. It requires large quantities
of cleaning air and energy
to heat the cleaning air
to bed temperature.
-------
The moving bed design also has some limiting operating charac-
teristics. Particle reentrainment caused by the relative motion
of the granules limits the granule flow rate and affects the over-
all collection efficiency. Erosion of the retaining grids, louvers,
and transport system components may be a problem, especially in
high temperature and pressure systems. The collected dust par-
ticles cannot form a filter cake so that the operating efficiency
will be essentially that of a clean bed.
No suitable HTP mechanical granule recirculation system and
granule/dust separation system were identified in the literature.
CPC used a pneumatic method. Granule transport and granule/dust
separation by the pneumatic method add significantly to the opera-
ting cost because large quantities of compressed air are required.
In order to maintain the high gas temperature at the gas turbine
inlet, it is necessary to heat the transport air to bed tempera-
ture to minimize the heat energy loss from the recirculating gran-
ules. This can be an important factor in the operating cost.
It may be possible to resolve most of these problems through
further development and testing. Performance data at high tem-
peratures and high pressures will be important in identifying
the most serious operational problems.
Intermittently Moving Bed -
In the late 1950s, Squires modified the continuously moving
bed design to obtain a fixed bed device with an intermittent move-
ment of granular solids. The bed is stationary during filtration.
The accumulated filter cake and the surface layer of granules
are ejected from the panel by a sharp backwash pulse and fall to
the bottom of the filter vessel. The expelled granules are immed-
iately replaced by downward movement of fresh granules from the
overhead hoppers.
The principal advantage of this type of bed structure is the
capability for external granule/dust separation with minimum dis-
turbance to the rooting cake. A rooting cake is the foundation
for the formation of a surface cake. After cleaning, the surface
cake is formed readily without disturbing the rooting cake and
filtration efficiencies can be much higher.
10
-------
The intermittently moving bed also has the advantage that
granule cleaning is off-line and potentially more effective.
The major disadvantage is that the gas capacity is lower than
for other granular bed filter designs and this results in high
capital costs for a given installation. The face velocity is
about one third the velocity used in the fixed bed and continuously
moving bed GBFs. Thus, more filtration area is required.
Bed plugging also can be a problem if the surface cake is not
formed properly. Erosion of the retaining grids, louvers, and
other components may be a problem.
For HTP applications, Squire's design has another disadvantage.
It requires large quantities of compressed air for cleaning. Com-
pressed air requirements could be as much as II of the gas treated.
The cleaning air is mixed with the gas being treated. In order
to maintain the HTP conditions at gas turbine inlets, it is neces-
sary to preheat the cleaning air. Cost to heat the cleaning air
can be an important factor in the operating cost.
Fixed Bed Filters -
Fixed bed filters operate in two modes; the filtration mode
and the cleaning mode. During filtration the bed is stationary.
The gas passes through the bed and collected particles are depo-
sited within the bed and on the bed surface. During cleaning the
bed is isolated from the main flow and agitated mechanically or
pneumatically by a reverse flow of gas.
There are two fixed bed devices currently being developed;
the Rexnord gravel bed filter and the Ducon granular bed filter.
The Rexnord filter uses a rake-shaped stirring device to agitate
the bed during cleaning. This loosens the filter cake which is
then removed by a reverse flow of clean air.
The Ducon granular bed filter cleans the bed by a reverse
flow of gas which fluidizes the bed and elutriates the collected
particles.
The Rexnord filters have been used successfully to control
emissions from clinker coolers in the cement industry. They
operate in the range of 100 to 200°C and near atmospheric
11
-------
pressure. No Rexnord filters have been tested at high temperature
and pressure.
The Ducon filter was tested on the effluent from a fluid bed
catalytic cracking unit regenerator at an oil refinery. The gas
was at 370°C to 480°C and 1 to 1.5 atm. A collection efficiency
of 85 to 98% was obtained on dust with a mass median diameter of
35 urn and a geometric standard deviation of about 4.
Various high temperature and pressure designs of the Ducon
filter were tested at the Exxon miniplant (Hoke, et al., 1978).
A number of operating problems were encountered during these tests.
The lowest demonstrated particulate outlet concentration was
68.6 mg/Nm3 (0.03 gr/SCF) which was considered to be too large to
protect a gas turbine and borderline for meeting current emissions
regulations. The use of smaller filter granules could be expected
to improve efficiency. However, at times the filtration efficiency
was very poor and the outlet particulate concentrations were as high
as 700 to 1,200 mg/Nm3 (0.3 to 0.5 gr/SCF). It was also observed
that the efficiency decreased with time in the longer runs, drop-
ping from 90% initially to about 50% later in the run. Loss of
filter medium during blow back was another recurring problem.
Further attempts were made to use 50 mesh retaining screens but
they failed because of plugging. Additional tests made with 10
mesh screens also resulted in significant screen plugging.
A significant buildup of particles in the filter beds was
also observed amounting to about 30% of the weight of the filter
medium. A possible steady long term increase in filter pressure
drop may result because of this. However, no significant increase
in filter pressure drop was noted during any of the shakedown runs.
It was also observed that the particles were not only building
up in the beds but were uniformly mixed with the filter medium.
It is possible that the buildup and mixing of particles in the
bed could be responsible for the increase in the particle concen-
tration in the outlet gas with time.
Another potential problem with the current design was its
vulnerability to upsets. When upsets occurred, such as bed
12
-------
plugging or loss of filter medium, the operating problems caused
by such upsets required shutdown of the system. Another problem
which may be unique to the miniplant was the interaction of the
granular bed filter with the rest of the FBC system during the
blow back cycle. An increase in system pressure was noted during
blow back resulting in problems with the coal feed system which
is controlled by the differential pressure between the coal feed
vessel and combustor. This required modifications to the coal
feed control system to minimize the effects.
The Exxon's granular bed filter test program was suspended in
November, 1977. In all runs in which more than one outlet concen-
tration was measured, it was observed that the outlet concentration
increased with time. They were not able to demonstrate that the
current EPA emission standard (0.1 lb/106 Btu or 0.05 gr/SCF)
could be met for more than a few hours of operation. In no runs
was the anticipated new standard (0.03 lb/106 Btu or approximately
0.05 gr/SCF) satisfied.
ENGINEERING DESIGN EQUATIONS
The literature search provided information concerning
theoretical and empirical design and performance equations for
GBFs. Table 3 is a summary of the available equations. Most of
the theoretical work considers the flow through a packed granular
bed to be similar to the flow around single granules. This
assumption appears to be inadequate for describing the efficiency
and pressure drop of actual granular bed filters. These equations
do not adequately predict the collection efficiency of a GBF at
face velocities normally encountered in practice.
In this study, a performance model was developed which pre-
dicts collection efficiency for granular bed filters. The
model is based on the assumption that the GBF is equivalent
to a large number of impaction stages connected in series.
13
-------
TABLE 3. AVAILABLE EQUATIONS FOR THE PREDICTION OF
PARTICLE COLLECTION IN A GRANULAR BED
Investigator
Equation
Notes
Jackson and Calvert
(1968)
Paretsky et al.
(1971)
Miyamoto and Bohn
(1974)
Gebhart et al.
(1973)
Bohm and Jordan
(1976)
Pt = exp
= exp
"c z-il
. ' -c "J
[- I 1=1 2- nl
L 2 e d I
['
N
pe
-6.
Pt
36 U
18e
UG\ 4 f Z
' yG dc
Impaction only
Collection by
diffusion only,
Collection by
diffusion only
Collection by
diffusion.
Gravity settling
and impaction.
Continued
-------
TABLE 3. (continued)
Investigator
Equation
Notes
Goren
(1977)
Pt = exp
i (1-e)
2
/
\
1250 K
z-zs
Collection by impaction,
gravity settling, and
diffusion
Westinghouse
(Ciliberti, 1977)
Ptj = exp
-3.75
- 5.8
dc
D
Collection by impaction,
interception, and diffusion
Schmidt, et al .
(1978)
Ptd =
dc I NPe
2.038 N
1/6
5/H
Re
+ 1.45
+ 3.97 K + —
p
Semi-empirical equation,
collection by diffusion,
interception, impaction,
and gravity settling.
-------
The model is based on the collection of particles in a clean
granular bed. The collection efficiency should be higher than pre-
dicted if there is a significant filter cake on the bed surface
and within the bed. The presence of a surface cake, however, has
not been observed by investigators working with large-scale fil-
ters. The clean bed model predicts a conservative estimate of
the efficiency attainable by granular filtration and is a satis-
factory model for filters which operate primarily without the
presence of a filter cake.
This model has been used to predict the performance of indus-
trial GBF systems. The particle collection efficiency predicted
with this model compares favorably with available field data.
POTENTIAL FOR HTP APPLICATIONS
The feasibility of advanced energy processes, such as pres-
surized fluidized bed coal combustion, depends on the availability
of a very efficient HTP particulate cleanup device. The parti-
culate control equipment should be capable of operating at a gas
temperature up to 9SO°C and a gas pressure up to 20 atm.
The suitability of GBFs for controlling particulate emissions
from advanced energy processes is not limited by the gas tempera-
ture and pressure. By properly selecting adequate granules and
structural materials, the granular bed filters could be capable
of operating at any temperatures and pressures encountered in
advanced energy processes. The Ducon GBF, Combustion Power Com-
pany's moving bed filter, and the CCNY panel bed filter, all can
be designed to operate at HTP.
The use of GBFs for HTP applications is limited by the par-
ticulate removal efficiencies and operating difficulties. Required
efficiencies depend on the future emissions standards and on the
tolerance of gas turbines for fine particles. Turbine require-
ments are not well established at this time.
At this time granular bed filters have not been demonstrated
to be efficient enough to perform as the final cleanup stage in
high temperature and pressure gas cleanup systems.
16
-------
There are several methods that may be used to increase the
collection efficiency. One method is to use a deep bed of fine
granules. This is not a good approach as the pressure drop would
be very high. Other methods such as electrostatic augmentation
and cake filtration should be more effective.
Quantitative data on the costs of HTP granular bed filters
are difficult to find. However, we have completed cost comparisons
for various types of GBFs. The cost estimates are based on the
cleanup requirement for a combined cycle steam turbine-gas turbine
power plant.
The estimated capital costs of moving beds and intermittently
moving beds are about 141% and 247% higher than those of fixed
beds,respectively. With regard to operating cost (not including
depreciation), moving beds are about 7.4 times higher than fixed
beds and intermittently moving beds are 4 times higher than fixed
beds.
The Energy Conversion Alternatives Study (EGAS) reported
estimates that the difference in the overall cost of electricity
between a pressurized fluidized bed boiler without cleanup and a
conventional steam power plant with stack gas cleaning is about
14 mills/kWhe. Thus, a gas cleanup cost on the order of a few
mills/kWhe should allow sufficient economic advantages to warrant
continued development of the PFB boiler process.
Based on the preliminary cost analysis, all three of the GBF
systems appear to be economically competitive. However, at the
present stage of development, GBF performance is neither efficient
enough nor sufficiently reliable to satisfy HTP particulate con-
trol requirements. Therefore, relative cost estimates must be
considered highly speculative and serve mainly to indicate areas
where further development work might substantially reduce costs.
CONCLUSIONS
The principal objectives of this study were achieved and
the following conclusions may be drawn.
17
-------
Present Technology
1. GBFs of present design are used successfully to control
emissions from clinker coolers in the cement industry and on hog-
fuel boilers in the forest products industry. The particles
emitted from these sources are relatively large and present GBFs
are capable of cleaning the gas sufficiently to meet the emission
standards.
2. GBF collection efficiency may be increased by using smal-
ler granules and deeper beds. It may also be increased signifi-
cantly by imposing an electrostatic field on the bed and by
building a good filter cake on the bed surface.
3. Further research work is required to increase the reli-
ability of the bed cleaning methods and granular solids transport
systems.
Design Equations
1. None of the design equations reported in the literature
are adequate to predict the collection efficiency of a GBF at
face velocities normally encountered in the field. A performance
model is presented which predicts collection efficiency for GBFs.
The particle collection predicted by this model compares favorably
wtih available field data.
2. The pressure drop across a clean bed may be predicted by
Ergun's equation for a packed bed.
HTP Potential
1. Granular bed filters may be designed to operate in high
temperature and high pressure environments. The available GBF
designs show potential for cleaning the gas to meet the current
New Source Performance Standard if fine granules (<500 pm in dia-
meter) are used as bed material.
2. Limited research work has been conducted to demonstrate
the feasibility of simultaneous removal of particulates and gaseous
pollutants with GBFs. It has been shown that packed beds of dolo-
mite are capable of removing sulfur dioxide and particles simul-
taneously. Celatom MP-91 diatomaceous earth, Burgess No. 10
pigment and activated bauxite appear to be possible candidates
18
-------
as packed bed sorbents for removing alkali metal vapor from hot
combustion gases.
3. The capital and operating costs of HTP GBF systems appear
to be within the acceptable range. A combined-cycle power plant
with GBFs for HTP participate cleanup is economically competitive
when compared to a conventional power plant.
4. Although granular bed filters appear to have the potential
for controlling particulate at high temperature and pressure in an
economically acceptable manner, they are far from a proven, state-
of-the-art technology.
There are many operational problems and uncertainties which
need to be resolved before HTP granular bed filters can be con-
sidered sufficiently reliable for commercial application. These
problems include the need to:
a. Prevent particle seepage through the bed (during cleaning
or filtration.
b. Reduce temperature losses (especially during cleaning).
c. Improve the efficiency and reduce the cost of granule
regeneration and recirculation.
d. Prevent attrition of granules causing particle reentrain-
ment.
e. Prevent sintering of granules.
f. Prevent plugging of retaining grids.
g. Reduce pressure drop across the bed.
h. Improve primary and overall fine particle collection
efficiency.
19
-------
SECTION 2
INTRODUCTION
High temperature and pressure gas streams are encountered
in the development of advanced energy processes such as coal
gasification and pressurized fluidized bed combustion. To in-
crease the thermal efficiency of these processes it has been
proposed that a combined gas turbine and steam turbine cycle be
used for the generation of power. However, the presence of
particulate matter in the gas at the gas turbine can damage
the turbine blades and render them inoperative. Thus it is
necessary to remove the particulate matter to a level which
will meet the turbine requirement for gas cleanliness with
minimal loss of gas temperature and pressure.
There are a number of particulate removal systems under
development that are capable of operating at high gas temperature
and pressure. The granular bed filter is one of these systems.
Names like "granular bed filters," "gravel bed filters,"
"panel bed filters," "sand filters," "moving bed filters," and
"loose surface filters" are used by various researchers to des-
cribe the type of air filtration equipment which consists of
a bed of graded sand or gravel. In this report the general
term "granular bed filters" is used and it is defined as any
filtration system comprised of a stationary or slowly moving
bed of separate, relatively close packed granules or particles
as the filtration medium. In order to prevent the collected
particulate matter from plugging the interstices between the
granules and causing excessive pressure drops, the device should
embody some means for either periodic or continuous removal
of the collected particles from the collecting surfaces. This
description then excludes fluidized or dispersed beds where
granular particles are kept in motion by the gas being treated.
It does include fixed bed or closely packed moving bed systems.
20
-------
This report reviews and evaluates the status and potential
of GBF technology for air pollution control with emphasis on high
temperature and high pressure applications. The principles of
operation, design performance of granular bed filters for removing
fine particles from gas streams, and the effects of physical
parameters such as granule diameter, bed depth and the design
of bed containment structures are discussed. Filtration mechan-
isms are described, as are several alternative procedures for
regenerating the granular bed filters.
Current practices in the application of granular bed filters
for air pollution control and existing granular bed filter systems
are critically reviewed and evaluated. Usage problems are iden-
tified.
The potential of granular bed filters for control of parti-
culate emissions from advanced energy processes such as fluidized
combustion and combined cycle power plants is evaluated. For
each potential application, the energy costs of high temperature
and high pressure particulate cleanup are compared with conven-
tional power plants.
21
-------
SECTION 3
LITERATURE REVIEW
PATENTS
Granular bed devices and processes for the removal of
particles from gas streams have been reported as far back as
the late 1800's. The following is a brief description of
principal patents issued since then.
Solvay (1889) patented a filter to remove dusts and vapors
from gases. It consisted of a cylindrical granular bed of
sand (or alternatively a fibrous bed of asbestos) arranged in
layers of increasing fineness upward from a foundation bed of
coarse gravel or pebbles. A steam jacket was added to the
vessel if a condensible vapor was to be removed. An internal
rake or scraper on a central vertical shaft provided cleaning
during operation. This pioneering patent contained many of
the ideas later incorporated into the design of large coke beds
for the removal of sulfuric acid mists and large sand beds for
the removal of radioactive particles.
The early decades of the twentieth century produced a
series of diverse patents involving granular filters. Fiechter
(1919) patented a device in which the gaseous medium to be
cleaned is introduced at the top of a rotating disc of perforated
metal (or a screen mesh) covered with a layer of sand or other
granular filter material. The gas is purified as it is passed
downward through the sand layer and is drawn off at the bottom
by a suction fan. Removal and purification of filter media
may be done continuously or intermittently using a revolving
screw conveyor, mounted over the disc, and a vertically adjus-
table scraper bar to spread the cleaned filter medium to the
desired layer thickness. The method of cleaning the spent
sand or filter medium is not specified.
22
-------
Another patent by Fiechter (1922) involved the use of a
moving sieve on an endless belt, carrying a horizontal layer of
sand through which a gaseous medium can be filtered downward
under suction or pressure. A pair of inner guide walls prevent
sand from slipping off the belt, which is slightly lower at one
end, allowing spent sand to trickle into a cleaning device and
be returned via an elevator to the hopper above the opposite or
higher end of the belt.
Klarding (1921) patented a filter device for constantly
purifying hot blast-furnace and generator gases, containing
large quantities of dust, without reduction in the temperature
of the gases. Granular filtering material flows continuously
from an upper hopper, downward through a main filter chamber,
through a lower funnel onto a vibrating sieve for dust removal
and then onto a conveyor that returns the clean filter material
to the upper feed hopper. Dusty gases are introduced at the
side above the main filter bed, flow downward through the bed,
and exit at the opposite side. Part of the cleaned gas is
directed against the dust-laden filter material as it is
shaken on the sieve to aid in removing the dust, which falls
through the sieve.
Nordstrom (1922, 1924) devised an improved means of sepa-
rating dust, smoke, and the like from gases in a cement-burning
process, a process for manufacturing chloride of lime, and the
copper smelting process. The gases are passed through a gran-
ular filtering material in a filter tower. A separate current
of atmospheric air moves dust from the filtering material as
it falls through a step-like bottom chamber; cleaned filter
material is conveyed back to the top of the filter; the removed
dusty matter may be returned to the original process or used
for another purpose. The filter tower, located between a
furnace and a chimney, consists of two concentric perforated
walls with the filter material contained between them. Gases
23
-------
from the furnace enter the inner chamber and pass outward
through the filter wall to an outer chamber connected to the
chimney and are discharged to the atmosphere in a purified
state.
Thomson and Nisbet (1924) invented a filter for cleaning
dust-laden gases from a blast furnace by allowing the gases to
be drawn horizontally through a vertical downward moving screen
of suitable ballast material. Ballast is continuously fed into
a V-shaped hopper at top, slips downward over metal slots
arranged in louver fashion, and is conveyed by a worm extractor
down a chute to a sloping metal screen at the bottom. The screen
can be agitated to facilitate separation of dust from ballast,
which is then conveyed to the top hopper by an elevator mechanism.
Suggested ballast material includes granulated or coarsely
powdered quartz, flint, or metallic fragments.
Lynch (1930) patented a filter designed to handle large
volumes of air or gas at high temperatures. It consists of a
thick bed of granular filter material falling downward into
piles in separate chambers. Gas flows through the bed at a
slow rate, not exceeding 3 m/s (10 ft/sec), and discharges in
a direction approximately opposite to the direction of filter
material flow after passing downward and then upward through
each chamber. Filter material is continuously cleaned, being
carried by conveyor to a rattler or other device for dust
removal, and then is hoisted back to the feed, chutes by an
elevator with bucket conveyor.
Lynch (1936) described a granular filter consisting of a
bed of gravel 1.2 to 2.5 cm (0.5 to 1 inch) in diameter that is
continuously withdrawn from the bottom of the filter, passed
over a screen for dust removal, and returned to the top of the
bed. Superficial gas velocity was approximately 91 cm/s (3 ft/s)
for beds 30 to 122 cm (1 to 4 ft) deep and pressure drop was
about 2.5 cm (1 inch) W.C. Units of steel, high chromium steel,
and brick have been used to filter gases having temperatures up
to 454'C (850'F), 816°C (1,500°F), and 1,093°C (2,000'F), respec-
tively.
24
-------
Fournier (1936) designed an apparatus for filtering gases
by means of a filtering material such as sand falling over
horizontal slats that may be vibrated on a combined system of
slots and sieves. The gas passes transversely through the
layer of filtering material and between the slots of each series
Hammer vibrators are suggested as a means of increasing the
filter surface by facilitating the flow of filter material
and partially cleaning it. An endless chain of buckets or
rakes continuously feeds clean sand into the upper hoppers and
transports soiled sand flows by gravity down a sinuous channel
against an upflow of cleaning gas that removes the dust to a
cyclone dust separator.
Berry and Fournier (1939) presented a more limited version
of the above as a German patent. Dust, soot, etc. are removed
from gases and vapors by passing the gas transversely through
a granular filter of material falling in piles with natural
angles of repose over a series of horizontal slats. The angle
of the slats may be varied, and filter material may be removed
at the bottom of the apparatus, cleaned, and replaced at
the top.
Carney (1944) devised an apparatus for separating carbon-
black dust entrained in a stream of gas or air by passing the
gas upward through a bed of carbon-black granules contained
in a rotating cylinder. A spiral conduit, connected at the
bottom of the cylinder and wrapped around it, rotates with the
cylinder and lifts the granules to the top of the cylinder
while the carbon dust is agglomerated to the granules
Mercier and Ehlinger (1950) developed a filter using
sand or other granular material to remove dust from hot gases
issuing from a boiler firebox or from a boiler heated by gases
under pressure. The filter is designed to handle gases of
any temperature or pressure. In principle, a sand of suitable
size and quality is arranged to provide small volumes, small
depth of mass, and considerable surface area so that gases may
pass through the filter without excessive pressure loss. Fine
25
-------
sand is held by walls formed by cone frustrums vertically
aligned and concentrically disposed approximately opposite one
another with conical surfaces tapering in opposite directions.
For coarse sand (> 4 mm), vertical layers are held between two
grids or between two perforated cylindrical metal sheets. The
filter unit is radially divided into cells and is rotated.
Gases to be cleaned enter at the top of a container de-
signed to withstand a high pressure (such as 150 kg/cm ), zig-
zag through the sand filter, and exit at the bottom during one-
half revolution of the filter. During the other half-revolution,
sand and dust are emptied down pipe to a rotating screen.
Dust falls through the screen while the sand passes into a
screw conveyor or similar device to be raised to the top of
the unit for recycling. Continuous cleaning and reuse of the
sand permits uninterrupted operation without loss of the heat
the sand has acquir-ed through contact with the hot gases.
Veron (1951) designed a sand filter for removing entrained
dust from gases produced by the combustion of pulverized coal
at high temperature and pressure and destined to feed a turbine.
The filter is contained in a cylindrical body having a domed
cap and conical bottom - a form suitable for high gas pressures.
An inner lining able to withstand high temperature forms a
cooling jacket, through which compressed air is circulated.
Gas enters at the bottom of the filter and passes upward through
stepped tiers of sand in multiple trays through a network of
interconnecting channels and pipes. Clean gas then flows
through a separate system of inner chambers between the trays,
exiting at the bottom of the filter on the opposite side of
the filter.
When the sand becomes heavily dust laden, as indicated by
an increase in pressure drop, slide valves are opened indi-
vidually, tier by tier, releasing the sand for cleaning by any
suitable method, and then the sand is returned to the top of the
filter. Sand was chosen as a readily available filter medium
that can stand high temperatures without damage or diminished
filtering power.
26
-------
Various designs of granular bed filters for removal of
aerosols from industrial air and gas streams continued to appear
in the 1950's and 1960's. Among those are Dorfan Impingo Filter
(U.S. Patent 2,604,187) Squires (U.S. Patent 3,296,775), Berz
and Berz (U.S. Patent 3,090,187), Kalen (U.S. Patent 3,798,882)
and Zenz (U.S. Patent 3,410,055; 3,880,508). All these designs
were tested at least in pilot scale and some were introduced
commercially. These designs will be discussed in a later section.
GBF STUDIES
Low Temperature Filtration
Among the earliest of granular bed filter tests were those
carried out in 1948-1949 at the General Electric Company, Han-
ford Works. The granular bed filters were used to remove radio-
active particles from air ventilation systems.
The results of the tests were reported by Lapple (1948).
The test chambers consisted of vertical cylinders, 25 to 30 cm
(10 to 12 inches) in diameter and 91 to 153 cm (3 to 5 ft) long. The
filter sand was supported on a 8 mesh screen welded to the sides
of the tank upon which was placed successively finer grades of
sand. Gas from the main ventilation system was passed through
the bed in an upflow manner for the tests.
Total concentration of suspended particles at the inlet
was on the order of 2.46 mg/m3 (0.001 gr/ft3) of which less than
0.025 mg/m3 was radioactive. Particle diameter was 0.5 to 2.0 ym,
Effects of bed depth, granule diameter, types of sand, and gas
velocity on collection efficiency were studied. Collection effi-
ciencies were based on the decrease of measured radioactivity of
the gas passing through the bed rather than on a decrease of
total particle concentration.
Table 4 shows a summary of the test results. At low gas
velocities (<5 cm/s), laboratory tests indicated that collection
efficiency increased with increased bed depth and finer sand
grain size. Irregular grained sands gave better collection effi-
ciencies; increasing gas velocity lowered collection efficiency.
The following approximation was derived:
27
-------
TABLE 4. RESULTS OF HANFORD SAND FILTER TESTS
Effect of Depth of Sand
Depth of Sand (cm)
30.5
61
91.5
122
Collection Efficiency,
91.6
99.5
99.8
99.9
Effect of Sand Size
Sand Type
Hanford-16 + 20 mesh
Hanford-20+40 mesh
Hanford-16 + 20 mesh
Hanford-20 + 40 mesh
Ottawa-20+30 mesh
Ottawa-30 + 40 mesh
Ottawa-20+30 mesh
Ottawa-30 + 40 mesh
dc (cm)
0.1
0.056
0.1
0.056
0.071
0.048
0.071
0.048
(cm/s)
1.93
1.93
5.1
5.1
2.6
2.6
5.1
5.1
Collection
Efficiency,1
96.3
99.98
89.5
99.88
96.2
98.6
92.9
98.2
Effect of Type of Sand
Type of Sand
Hanford -20+40 mesh
A.G.S. Flint-30+40 mesh
Monterey Type G
Ean Claire Type G
Effect of Gas Velocity
ur (cm/s)
0.91
3.1
5.1
9.3
20.3
Collection Efficiency,
99.93
99.84
99.61
99.38
28
Collection Efficiency,
99.97
99.83
98.9
92.7
97
-------
Pt = exp
u 0.3 3 j 1.3 3
G C
(D
where Pt = overall particle penetration on an activity basis,
K = proportionality factor, dimensionless fraction
Z = bed depth, cm
UG = superficial gas velocity, cm/s
d = granule diameter, cm
On the basis of these preliminary tests, large scale sand
filters were designed and set up at Hanford with a capacity of
14.2 m3/s (30,000 CFM) each. The design face velocity was
3 cm/s (6 ft/min), and the granular layer consisted of a 61
cm (2 ft) depth of-20+40 mesh Hanford sand. Seven layers of
coarser sand were also used in the bed to aid in flow distribu-
tion. Overall pressure drops were in the range of 10.2 to 17.8
cm W.C. (4 to 7"W.C.), and the average collection efficiency was
about 99.7%. The potential lifetime of these sand beds was
estimated at about 5 years based on estimated accumulation of
particles.
Thomas and Yoder (1956 a, b) investigated the filtration
of aerosols through a column of lead shot. Their data are shown
in Figure 1. The experiments were conducted at low flow rates
and with large diameter granules so that the effects of inter-
ception and inertia must have been negligible; the efficiency
was very low. Collection was primarily by diffusion and
and sedimentation, the former prevailing in the ascending
branches of the curves and the later in descending. Hence,
collection efficiency decreased rapidly with an increase in
flow rate. Sedimentation caused the collection efficiency to
differ for upwards and downwards flow as shown in Figure 1.
Similar curves were obtained with columns of sand in which
collection efficiency diminished appreciably with increasing
flow rate and grain size; grains of irregular shape gave higher
efficiency than round ones.
29
-------
50
z
o
H
w
z
W
CU
10
J I
I I I I
1
0.5
Figure 1. Filtration of aerosols through bed packed with
lead shot (Thomas and Yoder data).
0 0.1 0.2 0.3 0.4 0.5 0.6 0.7 0.8 0.91.0
PARTICLE RADIUS,
D Flow up column 1.49 cm/s
• Flow down column 1.49 cm/s
O Flow up column 0.745 cm/s
A Flow down column 0.745 cm/s
Column height 89 cm
Column area 11.17 cm
Lead shot 0.15 cm dia.
30
-------
Zahradnik et al. (1970) carried out a study on the simultan-
eous removal of fly ash and sulfur dioxide using a granular bed.
The granular bed was a square cross-section shaft, total length
224 cm (8 ft); the top 214 cm (7 ft) served as the storage area,
the remaining 30 cm (1 ft) was the actual filtering section.
The cross-sectional area was 930 cm2 (1 ft2). The filtering
medium was a slowly moving bed of 0.16 to 0.32 cm (1/16 to 1/8 in)
diameter alkalized alumina particles with the solids rate con-
trolled by a vibratory feeder at the bottom. The gas, fly ash,
and sulfur dioxide flowed horizontally across the filter.
Data on the simultaneous removal runs are shown in Table 5.
The S02 removal efficiency was 100% as determined by standard
iodometric titration of inlet and outlet gas samples. The fil-
tration efficiency was determined by passing the gas through a
Cuno filter which collected all particles larger than 1 ym in
diameter. For the size distribution of fly ash listed in Table
6, the overall filtration efficiency was about 99%.
Taub (1970) studied the transient behavior of granular bed
filters while collecting dispersed fly ash. His results show
high efficiencies are possible with clean filters, but perfor-
mance deteriorates as the dust content of the filter increases.
When a clean bed is put into operation, particles are col-
lected in the interstices of the filter near the surface. As
the dust deposit builds up, the saturation zone or dust may
work its way through from the dirty to the clean side of the
filter. This results from the drag force exerted on the particle
deposits by the gas flow. As the deposit extends through
the bed, the bed is saturated with dust and reentrainment occurs
causing the collection efficiency to decrease. The result of
Taub's work (Figure 2) shows this trend. The collection effi-
ciency remains initially constant with time. As the saturation
zone extends through the bed, the collection efficiency declines.
The work of Taub shows that a 3.3 cm deep bed of 1,500 ym par-
ticles has a capacity of approximately 3,000 g/m2(0.61 lb/ft2)
before the collection efficiency begins to fall.
31
-------
TABLE 5. SUMMARY OF ZAHRADNIK ET AL.'S DATA
Experimental Parameter Run #6 Run #7
Gas flow rate, Nm3/min 0.83 0.84
Nominal space velocity, hr" 3,500 3,500
Sorbent rate, kg/hr 4.76 3.89
Temperature, °C 204 204
Inlet S02 mole fraction 0.006 0.006
Outlet sorbent loading, g/100 g 19 10
Fly ash loading, g/m3 6.45 7.09
Filtration efficiency, % 99.3 98.7
S02 removal, I 100 100
Pressure drop, cm W.C. 5.1 5.1
Sorbent diameter, cm 0.16 0.16
Size Distribution of Fly Ash
Particle diameter, urn % less than
104.8 100
82.2 75.5
71.6 50.0
60.8 40.0
54 23.6
39.6 17.0
32 13.9
27.4 11.0
32
-------
sz:
w
PL,
PL,
0.033 m deep bed, 1.49 mm spheres, 60 cm/s gas velocity
Figure 2 . Time vs efficiency - 0.033 m deep granular bed filtering
fly ash, (Taub's data) .
-------
TABLE 6. SLOPES AND INTERCEPTS OF FITTED LINES
Support
Screen
Support
Grid
Support
Superficial
Velocity
(cm/sec)
8.2
11.2
8.2
11.2
Particle
Diameter
(ym)
0.8
1.6
2.9
0.8
1.6
2.9
0.8
1.6
2.9
0.8
1.6
2.9
Slope
(Transfer
unit /cm)
0.050
0.076
0.27
0.054
0.080
0.306
0.042
0.058
0.336
0.044
0.074
0.362
Intercept
(Transfer units)
(Percent)
0.017 (1.7)
0.12 (11.0)
0.66 (48.0)
0.017 (1.7)
0.16 (15.0)
0.56 (43.0)
0.46 (37.0)
0.81 (55.0)
1.04 (65.0)
0.46 (37.0)
0.81 (55.0)
1.04 (65.0)
34
-------
Paretsky,et al. (1971) studied the filtration of dilute
aerosols by beds of sand. They studied a bed of-10+14 mesh
angular sand at superficial velocities between 0.3 and 80 cm/s
in a 5.1 cm diameter bed at bed heights of 3.7, 8.2, and 19.2
cm. Flow directions studied were vertically upward, vertically
downward, and horizontal. Figure 3 gives data obtained at a bed
height of 19.2 cm.
At superficial velocities less than about 20 cm/s, upward
flow exhibited higher penetration than downward flow, the differ-
ence in penetration being greater at lower velocity. Horizontal
gas flow through the sand bed resulted in penetrations between
those for upward and downward flows.
Figure 4 gives data obtained at 8.2 cm height for finer
sand (-20 + 30 mesh). The relationship between upward and downward
flow data is approximately the same for the coarse sand.
Gebhart, et al. (1973) published an extensive experimental
study on the collection of aerosol particles in packed beds con-
sisting of uniform glass spheres. Monodisperse aerosols were
produced by atomizing a diluted suspension of polystyrene parti-
cles and drying the spray with clean air. Aerosol particles had
diameters in the 0.1 to 2 \im range. Concentration measurements
in front and behind the packed bed were carried out with a Laser
Aerosol Spectrometer.
The bed filter consisted of a glass cylinder with an inside
diameter of 8 cm. The cylinder can be filled with glass beads
to a maximum height of 41 cm. The inlet and outlet of the filter
were funnel shaped to give uniform flow distribution over the
filter. Bead sizes investigated were 0.4, 0.16, 0.05 and 0.0185
cm in diameter. The bed was a horizontal bed and gas flow was
in a downward direction.
The particle penetrations are plotted against particle dia-
meter for four different bead sizes in Figures 5-8. The parameter
in each of the diagrams is the mean air velocity inside the fil-
ter; i.e., interstitial velocity. Interstitial gas velocity is
related to superficial gas velocity by:
35
-------
100
80
60
o
fr-
<
40
z
UJ
o.
20
UPWARD FLOW
Q DOWNWARD FLOW
-10+14 MESH
SAND BED THICKNESS
i i i i i
19.2 cm
0.1 0.2 0.4
2 4 7 10 20
SUPERFICIAL VELOCITY, cm/s
40
100
Figure 3. Penetration of 1.1 urn aerosol as a function of superficial gas
velocity: 10-14 mesh sand (Paretsky et al. data).
100
UPWARD FLOW
DOWNWARD FLOW
- 20+30 MESH
BED THICKNESS
0.1 0.2 0.4
2 4 7 10 20
SUPERFICIAL GAS VELOCITY, cm/s
100
Figure
4. Penetration of 1.1 urn aerosol as a function of superficial gas
velocity: 20-30 mesh sand (Paretsky et al. data).
36
-------
100
O
i—i
H
<
OS
H
PJ
2
UJ
20 L
10
0.1
Figure 5.
0.2
0.5
dp,
Eixperimental results obtained by Gebhart et al : Penetration
versus particle diameter for different flow velocities
and granule size.
;s
o
I—I
H
E-
U4
Z
w
D-,
Z=20.5 cm '0.72
0.1
0.2
d ,
Figure 6. Experimental results obtained by C^
et al: Penetration versus particle diameter
for different flow velocities and granule
size.
37
-------
100
§
1-^
i
0.1
0.1
Figure 7.
0.2
dp,
Experimental results obtained by Gebhart et al: Penetration
versus particle diameter for different flow velocities and
granule size.
10
101
~ »•
g" 1(>-1
I"-
"• 10-'
10-"
0.1
0.72
0.2
0.5
dp,
Figure 8. Experimental results obtained by Gebhart
et al: Penetration versus particle diameter
for different flow velocities and granule
size.
38
-------
ur
"Gi ' ~ C2)
The porosity of the bed, e, was found to be 0.385 for all bead
sizes used.
The diffusion and the sedimentation branches of the pene-
tration curves are clearly distinguishable. In between there
exists a penetration maximum. The position of this maximum is
shifted to larger particles with decreasing bead diameter. The
exact position of maximum penetration is a characteristic feature
of the size and shape of the bed material.
There is a difference in measured penetrations for downward
and upward flow, as shown in Figures 9 and 10. This is an indi-
cation that gravity is important. Downward flow gives lower
penetration because the vectors of the flow velocity and settling
velocity are complimentary. For aerosols with diameter less than 0.3
ym, the gravity effect is negligible.
Knettig and Beeckmans (1974) studied the capture of mono-
disperse aerosol particles in the size range 0.8-2.9 ym in a
screen-supported and in a grid-supported fixed bed of 425 ym
glass beads. Monodispersed aerosol particles (1:2 weight ratio
of uranine and methylene blue) were obtained from a spinning
disk aerosol generator. The charges on the aerosol particles
were neutralized by a radioactive source in the aerosol genera-
tor. The aerosol particles were sampled isokinetically upstream
and downstream of the test bed, and were analyzed quantitatively
by fluorometry. All experiments were performed at ambient tem-
perature and pressure. The bed was a horizontal bed with a dia-
meter of 12.7 cm. The gas flow was in an upward direction.
Figures 11 and 12 are experimental results. Figure 11 shows
collection efficiency, expressed in transfer units (number of
transfer units, NTU = -InPt), plotted against bed height, for
the bed supported on a screen, at a superficial velocity of 8.2
cm/s. The screen was a 100-mesh screen (open area 30.3%). Figure
12 shows a similar plot for a perforated aluminum plate, 1.6 mm
thick. The plate had 144 holes spaced at 10 mm from center to
39
-------
lo
100
*. 50
2
ta
O.
10
Figure 9
I I I I I I
0.727 urn
0 0.5 1 1-5
SUPERSTITIAL GAS VELOCITY, cm/s
Gravity effect in granular beds measured
with an upward and downward directed aerosol
stream (Gebhart et al. data).
o
I—I
t-
tu
D.
10
10
10
0.1
UP
0.05 cm
2.87 cm/s
41 cm
0.5
PARTICLE DIAMETER, pm
Figure 10. Gravity effect in granular beds measured
with an upward and downward directed aerosol
stream (Gebhart et al. data).
-------
en
H
C/]
1
•** i
U Z
Z
w
UU
u.
LU
w
OS
d « 2.9
= 8.2 cm/s
5 10
BF.D HF.IGHT, cm
Figure 11 . Capture efficiency (transfer units)
versus bed height for the grid supported
fixed bed (Knettig and Beeckmans data).
CO
{-
t—(
Z
=3
Of
UJ
u.
10
u
I—(
u.
O!
o:
d = 0.8 pm
0 ,
o 5 in
BI:D nr.iniiT, cm
Figure 12 . Capture efficiency (transfer units)
versus bed height for the screen
supported fixed bed (Knettig 6 Beeckman)
-------
center. The diameter of the holes was 0.794 ram, resulting in an
open area of 0.5691.
All of the curves gave a linear relationship between collec-
tion efficiency, expressed in transfer units, and bed height.
Straight lines were fitted to the data by least squares. Table 6
shows the results for other superficial gas velocities. A linear
relationship implies that collection efficiency per unit volume of
bed is independent of bed height. Impaction appears to have been
the primary collection mechanism in the body of the bed because
collection efficiency per unit volume of bed increased with both
superficial gas velocity and aerosol particle size.
Substantial aerosol capture occurred at the bed supports,
especially the grid support. This is because each grid hole
acted as a jet and caused the aerosol particles to be collected
on the glass beads by inertial impaction. However, the slopes
of the least square fitted lines for screen-supported and grid-
supported beds are the same.
Miyamoto and Bohn (1975) studied the effect of particle
loading on collection efficiency of granular bed filters. They
used water-washed river gravel for granules and ammonium chloride
fume for particles. The particle diameter was 0.1 to 3 ym.
The relationship between granular bed collection efficiency
and particle loading on the filter is presented for different
gravel diameter (Figure 13), bed depth (Figure 14) and superficial
gas velocity (Figure 15). Particle loading is defined as the
weight of particles collected per unit bed area. From Figures
13 through 15 the following observations may be made:
1. The collection efficiency increased with increasing
particle loading. The collected particles decrease bed porosity
and thereby increase filtering efficiency.
2. Smaller granule diameters result in higher collection
efficiency and a sharper increase of collection efficiency with
particle loading.
3. Thicker gravel layers had higher initial collection
efficiencies but little influence on collection efficiency at
42
-------
1.5mm
3.0mm
5. 2mm
7. 2mm
0 0.1 0.2
PARTICULATE LOAD, kg/m2
Figure 13.
Effect of mean gravel diameter
and particulate load on
granular bed collection
efficiency (Miyamoto and
Bonn's data).
U
tu
I—I
U
PU
E- U.
a
J
o
U
Z=10cm
=3 Ocm
= 5 Ocm
=7 Ocm
=90cm
Figure 14.
Effect of gravel layer thickne;
and particulate load on granuls
bed collection efficiency
(Miyamoto and Bohn's data).
PARTICULATE LOAD, kg/m"
=2.0 cm/s
=3.0 cm/s
=4.0 cm/s
"0 0.1 0.2
PARTICULATE LOAD, kg/m2
Figure 15.
Effect of superficial gas
velocity and particulate
load on granular bed
collection efficiency
(Miyamoto and Bohn's data).
43
-------
higher particle loading because the collection is preferentially
near the surface where collected particles tend to form a cake.
4. Higher superficial gas velocities markedly decreased the
rate of increase in efficiency with particle loading, probably
because filter cake is less stable at higher velocities.
Figueroa (1974) and Figueroa and Licht (1976) conducted an
experimental program to determine the aerosol filtration effi-
ciency of a bed of granular solids. The bed was a 10.2 cm (4
inch) I.D. column packed with plastic beads and sand to various
depths. The bed support was made of 325 mesh stainless steel
screen.
Two different size fractions of plastic beads (polyacrylo-
nitrile) and one grade of sand were used as granular bed solids.
The properties of the granular materials are listed in Table 7.
Two test aerosols were used: monodispersed methylene blue
(MB) particles 1 ym and 2 ym in diameter and monodispersed poly-
styrene latex microspheres 0.50, 1.10, and 2.02 ym in diameter.
Aerosol number concentrations before and after the bed were
measured with an optical counter.
Bed penetration was measured as a function of aerosol par-
ticle diameter, granule diameter, bed depth, direction of flow,
and superficial gas velocity. Figures 16 through 19 summarize
their data. For the low gas velocity range (20 cm/s), they
observed, for a fixed bed, that particle collection efficiency
always increases with decreasing granule diameter and increasing
bed depth. Downward flow gives higher collection efficiency than
upward flow. The collection efficiency of a granular material
with inherent electrostatic charging properties (e.g., plastic
beads) is higher than for a granular material without those proper-
ties (e.g., sand) .
i
Westinghouse Research Laboratories (Ciliberti, 1977) carried
out an experimental program to investigate efficiency and opera-
bility of granular bed filters. The experimental bed was designed
with a 233 cm (0.25 ft) cross section. The bed was cleaned
by fluidization. A distributor consisting of several independent
drilled tubes was used to control the fluidizing gas flow. The
44
-------
TABLE 7. EXPERIMENTALLY DETERMINED PROPERTIES OF THE GRANULAR MATERIALS
PROPERTY
Number
Median
Diameter
Sauter
Diameter
Geometric
Standard
Deviation
Solid
Granule
Density
Bulk
Density
Porosity or
Voidage
Fraction
(Figueroa, 1974)
SYMBOL UNITS
cN
3 2
g
micrometers
micrometers
micrometers
g/cm3
g/cm3
POLYACRYLONITRILE
495
1.10
0.67
0. 39
305
1.12
0.66
0.38
SAND
-25 + 40 mesh -40 + 60 mesh -25 + 40 mesh
680
514
1.13
339
1.23
702
1.12
3.10
1.85
0.41
-------
1.0
0.1
10'
|io-J
10"
10
d • 420 to 710 um PLASTIC BEADS
Z • 3 on
2.0 vm
Lt i i I A i i I i i i I i i i 0
12
10
3=
e
Q
£
0 4 8 12 16 20
SUPERFICIAL GAS VELOCITY, cm/s
Figure 16. Penetration tests on 420 to 710 micrometers plastic
beads by monodispersed polystyrene latex aerosol
(downflow) (Figueroa and Licht's data).
o
o
Q!
1.0
0.1
10"
10"
£ 10
10"
10"
d - 420 to 710 \im PLASTIC
d « 250 to 420 vim PLASTIC
c BEADS
9 cm
d - 0.5 urn PSL
I
4 8 12 16
SUPERFICIAL GAS VELOCITY, cra/s
20
Figure 17. Effect of bed height and bed granule (bead) size
on the downflow penetration of 0.5 micrometers
polystyrene latex aerosol particles on plastic
beads (Figueroa and Licht's data).
-------
m d - -25 +40 MESH SAND
10
0 4 8 12 16
SUPERFICIAL GAS VELOCITY, cm/s
Figure 18. Comparison of downflow penetration of polystyrene
latex aerosol particle on 420 to 710 micrometers
plastic beads and -25 +40 mesh sand granules
(Figueroa and Licht's data).
1.0
0.1
10"
o
H
CH
t-
5
10
10
I'M
d = 420 to 710 ym
PLASTIC BEADS •
- UPWARD GAS
F
^DOWNWARD
GAS FLOW
d = 0.5 m PSL
Z = 3 on
4 8 12 16
SUPERFICIAL GAS VELOCITY, cm/s
20
Figure 19 .
F.ffect of bed granule (bead) size and flow
direction on the penetration of 0.5 micrometer
polystyrene latex aerosol particles on plastic
beads. Bed weight = 150 grams, bed height = 3 cm
(Figueroa and Licht's data).
-------
arrangement of the bed is illustrated in Figure 20.
Bed depth ranged from 7.6 cm to 15.2 cm (3 to 6 in.) based
on requirements to minimize pressure drop while maintaining a
bed that can be successfully fluidized for cleanup. Granule
size was in the range of 15 to 30 mesh to achieve efficient dust
collection with small bed depth and reasonable pressure drop.
Superficial gas velocity ranged from 9 to 30 cm/s. The test
dust was finely ground (<10 ym) limestone dispersed in a high
velocity air jet. The layout of the experimental equipment is
shown in Figure 21.
Test results are summarized in Tables 8 through 12. Particle
size and mass data were taken with cascade impactors. Collection
efficiency for submicron particles was high. The grade effi-
ciency curves were flat and showed high collection efficiency
for all particle sizes. This was consistent with the observed
formation of a filter cake.
To investigate dust accumulation in the bed, Westinghouse
made a series of five runs. The initial clean bed material was
sampled, then after five consecutive runs bed samples were taken
at four levels through the bed.
The clean bed contained 0.5 wt I fine dust. After five
runs the dust content of the bed was 1.3 wt $. At the end of the
filtration cycle, the dust level at the bed surface was 10 wt %.
However, at levels 2.5 cm below the surface and greater the dust
level was uniform at 1.3 wt %.
A second test over ten cycles showed dust accumulation of
1.0 to 1.8 wt % in the bed.
Another filter, 0.65 m2 (7 ft2), was tested at temperature
of 1,100°F at atmospheric pressure. Because of operating pro-
blems tests were discontinued.
Combustion Power Company has conducted extensive cold flow
tests on a moving bed granular bed filter. These results will
be discussed later.
High Temperature Filtration
Dennis et al. (I960), in designing an incinerator for dis-
posal of low-level radioactive wastes from hospitals or biological
48
-------
FLUIDIZING GAS OUT
DIRTY GAS IN
FLUIDIZING
GAS IN
A - DEPTH OF SAND BED
B - DEPTH OF COARSE GRAVEL USED AS BED SUPPORT
FILTERED
GAS OUT
Figure 20. Schematic of granular bed filter
(Westinghouse setup).
49
-------
01
o
OUTLET
SAMPLE-
PORT
INLET SAMPLE
PORT
t
*
'»• • * • ;','.?•••'*•
DUST
HOPPER
COMPRESSED
AIR
Figure 21. Schematic of test equipment (Westinghouse).
-------
TABLE 8. WESTINGHOUSE GBF DATA
Series #2 Gas Velocity: 17.8 cm/s (+_ 10%)
Bed Material-20+30 mesh Ottawa sand
Bed Depth: 7.6 cm (3 in.)
Particle Size Inlet Dust Outlet Dust Collection
(Typical) (Typical) Efficiency,
mg/m3 mg/m3
0-0.3
0.3-0.45
0.45-0.75
0.75-1.5
1.5-2.3
2.3-3.3
3.3-5.0
5.0-8.0
8.0 +
8.0 +
30.6
71.8
120.6
132.4
59.0
10.5
12.3
10.5
10.0
14.1
0.93
1.81
2.6
1.5
0.3
0.25
0.25
0.05
0.1
0.1
96.9
97.5
97.8
98.9
99.4
97.6
98.0
99.5
99.0
99.2
Run No. Inlet Loading g/m3 Overall Efficiency,
2.7 0.28 98.8
2.8 0.47 98.3
2.9 2.5 97.9
2.10 1.5 97-5
51
-------
TABLE 9 . WESTINGHOUSE GBF DATA
Series #3
Gas Velocity: 10.2 cm/s (+_ 10%)
Bed Material-20+30 mesh Ottawa sand
Bed Depth: 7.6 cm
Particle Size
0-0.3
0.3-0.45
0.45-0.75
0.75-1.5
1.5-2.3
2.3-3.3
3.3-5.0
5.0-8.0
8.0 +
8.0 +
Inlet Dust
(Typical)
mg/m3
12.9
43.5
102.0
194.0
203.0
105.0
108.0
59.0
111.0
230.0
Outlet Dust
(Typical)
mg/m3
1.1
2.7
3.6
2.5
0.5
0.1
0.05
Collection
Efficiency, %
91.5
93.8
96.5
98.7
99.8
99.9
99.95
100.0
100.0
100.0
Run No.
3.11
3,12
3.13
3.14
3.15
Inlet Loading g/ms
0.6
1.16
0.72
0.72
0.78
Overall Efficiency,
97.2
99.0
99.2
97.7
99.4
(95.5)
(98.8)
(98.5)
(96.2)
(99.4)*
*Increased gas rate to 17.8 cm/s
52
-------
TABLE 10. WESTINGHOUSE GBF DATA
Series #5 Gas Velocity: 25.4 cm/s (+_ 10%)
Bed Material-20+30 mesh Ottawa sand
Bed Depth: 7.6 cm
Particle Size Inlet Dust Outlet Dust Collection
(ym)
0-0.3
0.3-0.45
0.45-0.75
0.75-1.5
1.5-2.3
2.3-3.3
3.3-5.0
5.0-8.0
8.0 +
8.0 +
Run No.
5-22
5-23
mg/m3
52
66
100
138
79
37
41
34
73
221
Inlet
mg/m3
0.1
0.14
0.2
Loading mg/m3
1.0
0.8
Efficiency, %
100.0
99.8
99.8
99.8
100.0
100.0
100.0
100.0
100.0
100.0
Overall Efficiency,
99.96
99.95
53
-------
TABLE 11. WESTINGHOUSE GBF DATA
Series #6
Gas Velocity: 25.4 cm/s (^ 10%)
Bed Material-16+20 mesh Ottawa sand
Bed Depth: 10.2 cm
Particle Size
Inlet Dust
mg/m3
Outlet Dust
mg/m3
Collection
Efficiency,
0-0.3
0.3-0.45
0.45-0.75
0.75-1.5
1.5-2.3
2.3-3.3
3.3-5.0
5.0-8.0
8.0
8.0
32
20
33
53
47
32
53
46
108
513
0.39
0.08
0.12
0.08
0.04
0.04
0.04
_ _
98.8
99.6
99.6
99.8
99.9
99.9
99.9
100.0
100.0
100.0
Run No.
6-24
6-25
6-26
Inlet Loading mg/m3
5.8
0.94
1.17
Overall Efficiency,
99.96 (99.92)
99.92 (99.8)
99.90 (99.7)
54
-------
TABLE 12. WESTINGHOUSE GBF DATA
Series #7
Gas Velocity: 17.8 cm/s (+_ 10%)
Bed Material-16+20 mesh sand
Bed Depth: 10.2 cm
Particle Size
Cum)
0-0.3
0.3-0.45
0.45-0.75
0.75-1.5
1.5-2.3
2.3-3.3
3.3-5.0
5-0-8.0
8.0+
8.0 +
Inlet Dust
Run 7-27
(mg/m3)
9.4
25.9
51.8
56.5
16.5
1.2
2.4
TOTAL 163.5
Inlet Dust
Run 7-28
(mg/m3)
11.7
42.4
73.0
88.3
41.2
7.1
2.4
1.2
266.0
Overall Efficiency
7-27
7-28
99.55%
99.58%
55
-------
laboratories, used a 20 cm (8 in.) layer of 0.64 cm (0.25 in.)
gravel to screen out coarse particles before the gas passed
through a 5.1 cm (2 in.) bed of slag wool. This filter unit,
with a 2,600 cm2 (2.8 ft2) filter area, was housed in half of
a 55-gal drum located approximately 244 cm (8 ft) downstream from
the incinerator. Gases exiting from the incinerator at 870 to
982°C (1,600 to l,800°F) were passed at negative pressure through
a water-cooled condenser so that filtration gas temperatures
were 93 to 427°C (200 to 800°F) with a pressure drop of about 2.5 cm
W.C. (1 in. W.C.). In one series of tests in which 408 kg
(900 Ib) of sawdust was burned, the pressure drop increased from
1.3 cm W.C. to 1.8 cm W.C. (0.5 to 0.7 in. W.C.) with a 90-98%
filter collection efficiency on a weight basis.
Strauss and Thring (1960) carried out studies on filtration
of submicron fumes from open hearth furnace gases using a granu-
lar bed. The bed was 5.1 cm (2 inches) in diameter and consisted
of 0.79 (5/16 in.) crushed high temperature insulating bricks.
Experiments were run with variety of bed thicknesses and gas flow
rates. Collection efficiency tests were carried out on cold and
preheated beds.
Table 13 shows data obtained from Strauss and Thring's granu-
lar bed filter study.
Collection efficiencies of 59.3 to 96.3% were obtained
with a bed depth of 25 to 26.7 cm (1 to 10.5 in.), average gas
velocities of 36.3 to 102.2 cm/s (1.2 to 3.4 ft/s), and maximum
pressure drops of 0.97 to 12.4 cm W.C. (0.38 to 4.9 in W.C.)
at gas temperatures from 230 to 520°C.
Further theoretical studies by Thring and Strauss (1963)
considered the effect of high temperature on particle collection
mechanisms. The controlling mechanism in these tests might be
inertial impaction, but they emphasized the importance of the
effect of inlet dust concentration on efficiency. It was main-
tained that dust particle agglomeration in the bed occurring
in the tortuous paths between the collecting granules plays
a highly significant role in collection. Increased mass flow
56
-------
TABLE 13. DATA OBTAINED FROM STRAUSS AND THRING GRANULAR BED FILTER STUDY
cn
Test Furnace
No. Operation
3
4
5
8
9
10
12
14
15
32
36
37
41
45
49
56
71
. 80
R*
R
C
C
R
0
R
A
T
M
M
R
R
R
F
R
R
R
Test Time
(min)
13
15
19
4.88
13.13
6.95
15.1
14.4
12.7
24.9
20.4
10.3
20.6
13.0
10.4
15.1
22.2
IS. 9
Gas Mass
Flowrate
kg/m2-hr
1,103
996
1,098
839
708
683
1,044
1,035
996
615
786
976
1,650
1,005
1,547
1,249
1,532
1,728
UG
cm/s
43.6
36.6
44.8
36.3
29.3
24.1
50.9
47.0
46.4
28.1
38.4
45.8
102.2
39.3
89.7
63.7
76.6
70.2
T~ Inlet Fume
or Concentration
L g/Nra3
230
230
275
290
270
237
345
303
317
305
352
300
520
225
460
370
360
455
1
4
0
6
3
6
0
0
0
0
0
1
1
0
0
0
1
4
.63
.14
.123
.26
.83
.24
.93
.54
.60
.127
.163
.24
.73
.595
.232
.573
.18
.94
Fe203in
Fume
Wt *
--
--
--
--
78
64
20
23
36
44.2
55.2
--
--
--
--
--
--
Z
(cm)
26.7
26.7
26.7
26.7
26.7
22.9
22.9
22.9
22.9
22.9
22.9
22.9
7.6
7.6
7.6
2.5
2.5
Z.5
Collection
Efficiency
Wt *
90.5
96.3
87.3
87.6
94.7
85.0
90.2
84.0
87.5
84.6
74.3
94.2
85.3
65.5
82.7
59.3
87.0
88.2
Collection Pressure
Efficiency of Drop
FezOj, Wt t cm W.C.
5.6
5.6
6.6
8.9
10.2
87.3 11.9
96.0 5.6
97.1 . 6.1
98.1 6.4
85.7 12.4
74.1 9.1
9.9
5.3
1.8
6.0
0.97
2.2
1.3
* These letters refer to the phase of the steelmaking cycle in progress at the time of the test.
0 = Oxygen lancing M " Melting
R - Refining c = Charging
-------
rate of gas and increased temperature also increased collection,
but to a lesser extent. When the temperature of the gases
differs greatly from that of bed, thermal precipitation plays
a role, but other mechanisms are more important.
Goldman (1964) reported that gravel bed filters have been
used for several years in Germany as large-pore filters that are
wear resistant in high temperature applications (to 350°C) . The
theory of the collection process is given briefly. Prior to
adoption of these filters, tests were made with various dusts,
including coke dust in the off-gases from a coke-drying operation,
phosphate dust, dust in the fumes from a carbide furnace, and
dust in the waste gases from a mixture of phosphorescents.
Gases contained 0.5 to 3 g/Nm3 of dust; the outlet gases con-
tained 10 to 95 mg/Nm3.
Dust removal was in the range of 93 to 97%, with pressure drops
of 11 to 20 cm W.C. and flow rates of 4,000 to 7,000 m3/hr. When
the pressure drop became too high, the gravel was washed and
the clean gravel returned to the bed for reuse.
The U.S. Bureau of Mines, Morgantown Coal Research Center
tested a GBF of Squires' design at high temperature (540°C,
1,000°F) (Wu, 1977). The bed is a vertical layer of sand held
in place by louvered walls. The filtering surface of the bed
was 7.6 cm (3 inches) wide and 30.5 cm (12 inches) tall. Several
grades of sand were used as bed material.
The test dust was derived directly from a boiler furnace.
Pulverized coal with 70% passing through a 200 mesh sieve was
burned. The coal was fed by a screw feeder driven by a variable
speed motor into a combustor with a capacity of 2.3 kg/hr (5 lb/
hr). The temperature inside the combustor was about l,093°-to
1,315°C (2,000 to 2,400°F). Natural gas was also burned during
the coal combustion in order to insure complete combustion
because of a coal feed rate problem.
At a face velocity of 10 cm/s (20 ft/min), the overall
collection efficiency was 98%. No particle size distribution
data were given.
58
-------
The City College of New York tested the same design of the
GBF at room temperature. Redispersed coal fly ash was used as the
test dust. The overall collection efficiency was 99.99%. The par-
ticle size distributions might not be the same for the two tests,
we cannot be sure whether the lower efficiency was a result of
operating at high temperature.
One observation regarding the difference in the filter cake
was recorded. In the Bureau of Mines' high temperature tests, no
filter cake over the sand surface was observed. At near room tem-
perature, a good surface filter cake was observed. Squires (Wu,
1977) attributed this difference in filter cake to the decrease
of adhesive and autohesive forces of fly ash at high temperatures.
Adhesion is defined as the interaction of particles with a solid
surface and autohesion is defined as the interaction of particles
among themselves.
A high temperature and pressure design of the Ducon granular
bed filter was tested at the miniplant of Exxon Research and•
Engineering Company (Hoke, et al., 1978). Exxon's experience with
this filter will be discussed in a later section.
GBF With Flux Forces
Anderson and Silverman (1957, 1958) reported on a study con-
ducted at the Harvard University Air Cleaning Laboratory to inves-
tigate electrostatic filtration in fixed and fluidized granular
beds. They observed that triboelectrification or friction char-
ging of fibrous and granular filter media can improve collection
efficiency with no increase in flow resistance. Similar results
have recently been reported by Figueroa and Licht (1976).
Fuchs and Kirsch (1965) conducted tests to determine the
effect of vapor condensation on the collection efficiency of gran-
ular beds. Monodispersed aerosols of selenium in nitrogen (par-
ticle diameter of 0.2 and 0.4 ym) were passed through a column
of silica gel, both directly and upon addition of some ether vapor.
It was found that with the addition of ether vapor, the collection
efficiency of the granular bed increased. Condensation of the
vapor and deposition of the particles occur simultaneously in the
59
-------
bed. The vapor molecules diffusing towards the granules sweep
the particles along with them.
A study of the collection of submicron aerosol particles in
an electrified granular bed was carried out by Research-Cottrell,
Inc., Bound Brook, New Jersey. One such device is described in
U.S. Patent No. 2,990,912, July 4, 1961. Collection of particles
occurred in a bed of 3 to 6 mm diameter glass beads held between
two screens maintained at different electrical potentials (Figure
22) . The aerosol particles were electrically charged upstream
from the bed. Data taken on collection of a 0.5 to 0.7 ym dia-
meter methylene blue particles are shown in Figure 23. It is noted
that a 99.5% efficiency was obtained in a. bed under 5.1 cm (2 in.)
in thickness with a pressure drop of 0.5 cm W.C. (0.2 in W.C.)
and a a face velocity of 61 cm/s (2 ft/s).
Sharapov (1975) tested a high gradient magnetic filter with
a bed of 8 mm steel balls for the removal of dust from open-
hearth furnace. The capacity of the filter was 60 m3/min. At
the optimum .voltage of 80 to 120 kV/m, the collection efficiency
was 80 to 901, and the energy consumption was 0.05 kWh/1,000 m3
of gas. Without the magnetic field the collection efficiency was
25 to 30%.
Gaseous Pollutants
Only limited work has been done in this area. Zahradnik et
al. (1970) and Squires and Graft (1971) have demonstrated the fea-
sibility of simultaneous removal of fly ash and S02 from gas
streams using a granular bed filter packed with alkalized alumina
and half calcined dolomite. Swift et al. (1977) are studying the
removal of alkali metal (sodium and potassium) compounds from the
combustion gases at high temperature with a packed bed. NaCl is
passed through beds packed with alundum, Celatom MP-91 diatomaceous
earth, Burgess No. 10 pigment (kaoline clay), attapulgus clay
(magnesium aluminum silicate), and activated bauxite. The tem-
peratures of the vapor and bed are maintained at about 870°C. The
gas velocity passing through the bed is 7.5 cm/s (3 in./s). Pre-
liminary test results revealed that Celatom MP-91 diatomaceous
60
-------
o\
Electrified
Beads
Charging Zone -=.
Figure 22 . Electrified packed bed.
-------
60
W
_J
u
t— (
0
Di
U
t
w
I— 1
u
I— I
tL,
UH
o
w
pa
(X
80
90
95
99
99.5
9-9.9
T 1 1 1 1
FACE VELOCITY =2.44 m/s (8.0 ft/s)-
.3)
1.22 m/s
(4.0 ft/s)
12.3).
= 0.61 m/s (2.0 ft/s)
(0.46)
fO.51.
NOTE
EFFICIENCIES DO NOT INCLUDE CHARGING SECTION
COLLECTION OF APPROX. 50%
FIGURES IN PARENTHESIS () DENOTE PRESSURE
DROPS (cm W.C.) ACROSS BED
I
I
1
I
0.6 0.8 1.0 1.2 1.4 1.6 1.8
PACKED BED THICKNESS, 6.0 mm SPHERES (INCHES)
Figure 23. Performance characteristics of electrostatically
augmented packed bed (Research-Cottrell data).
62
-------
earth, Burgess No. 10 pigment, and activated bauxite are promising
candidates as hot packed bed sorbents to remove alkali vapor from
hot coal combustion gases. Among the three, activated bauxite,
which is thermally treated high-alumina-content natural bauxite
ore, is the most active for removing NaCl vapor under the test
conditions.
THEORY
Particulate Collection
The primary mechanisms for particulate collection in a bed
of granular solids are:
1. Inertial impaction
2. Flow-line interception
3. Diffusional collection
4. Gravity settling
Inertial Impaction -
The inertial impaction mechanism is based on the inertial
force of the particle. The inertial force tends to move parti-
cles across the gas flow lines toward the collecting surface in
regions where the flow is diverging upstream of a fixed boundary
surface. Inertial forces increase with particle velocity and mass.
Diffusion -
Diffusion of the aerosol particle is based on the theory of
Brownian motion. The diffusion mechanism becomes increasingly
effective with decreasing particle diameter.
Gravity Settling -
The gravitational force can be significant for collection
of particles as small as 0.4 ym in diameter. This effect has
been demonstrated experimentally by the decrease in penetration
observed for downward flow compared with upward flow in sand
beds (Thomas and Yoder, 1956b). A gravitational force in the
direction of the bulk flow will tend to increase collection due
to settling across flow lines.
63
-------
Interception -
Interception is the mechanism whereby particles are collected
on surfaces while following the gas streamlines. Collection by
this mechanism usually is negligible for a clean granular bed.
However, during a filtration cycle, particles will deposit in the
interstices of the bed to form an internal cake and on the surface
to form a surface cake. As the cake builds up the bed porosity
and flow channels decrease and interception becomes an important
collection mechanism. When the flow channels are too small to
allow particles to pass, collection is ensured, being referred
to as complete interception or sieving.
In finely packed beds operated at low gas velocities, gravity
settling and diffusional deposition will predominate. The collec-
tion efficiency will be expected to decrease as the gas velocity
increases. Coarsely packed beds operating at higher velocities
(but still below fluidizing velocities) provide separation mainly
by inertial deposition and interception. The collection efficiency
will increase with velocity provided that the gas velocity is not
so high as to reentrain collected material.
The operation of granular bed filters is similar to fabric
filters even though granular bed filters have larger pore sizes
and deeper beds. Payatakes (1977) classified the filtration cycle
into four successive stages.
1. When the filter is new, particles deposit directly on the
granule surfaces. This is referred to as clean bed filtration.
The collection efficiency for this stage of filtration depends
primarily upon the granule size and the depth of the bed.
2. Particles deposit not only directly on granules but also
and preferentially on deposited particles, thereby forming par-
ticle dendrites.
3. The dendrites grow to the extent that they intermesh with
their neighbors forming a particulate coating around each granule
which is non-uniform in thickness.
4. If the granules of the bed are sufficiently small, particle
coatings of neighboring granules will bridge the gap to form an
internal cake. Lee (1975) called the internal cake a rooting cake
64
-------
and it is the foundation which supports the formation of a surface
cake. Once a surface cake is formed, filtration efficiency no
longer depends upon the depth of the granular bed but rather on
the thickness and structure of the surface cake. The cake filtra-
tion results in a much higher efficiency than the original granular
bed, and particle collection by sieving becomes a more important
collection mechanism.
Mathematical Models -
Currently there are several mathematical models available
for the prediction of particle collection in a granular bed. All
models are for the prediction of particle collection by clean
beds; i.e., stage 1 filtration. Stage 1 filtration is very brief
compared to the total filtration cycle.
Model by Jackson and Calvert - Jackson and Calvert (1966) and
Calvert (1968) have developed a theoretical relationship between
particle collection efficiency and packed bed operating parameters.
They assumed that the gas (and particle) flow through the bed may
be modeled by the flow through a series of semicircular channels
and that particles are collected by centrifugal force on the out-
side channel walls as gas and particles pass through the channels.
Their equation for predicting the particle penetration for a packed
bed is:
Ft, . erp I- C, -± K I (3)
where Ptj = penetration for particles with diameter "dp",
fraction
C: - empirical constant, dimensionless
Z = depth of the bed, cm
d = granule diameter, cm
K = inertial parameter, dimensionless
V C< S UG
9 HG dc
65
-------
d = particle diameter, cm
C1 = Cunningham slip factor, dimensionless
p = particle density, g/cm3
Up = superficial gas velocity, cm/s
yG = gas viscosity, g/cm-s
The empirical constant, "Cy, is a function of bed porosity,
channel width and granule diameter. It can be calculated by
the following formula:
C =
IT , .x
(4)
— v j j j —
where j = ratio of channel width to packing diameter, dimen-
sionless
e = bed porosity, dimensionless
For a packed bed with a granule diameter of 1.27 cm (0.5 in),
the empirical constant was found to be 21.4.
Model by Paretsky et al. - Paretsky et al. (1971) proposed
the following equation (based on Happel's "free surface model")
for particle penetration through a granular bed:
Ptd =
- 2 dc J
where Ptd = penetration of particles with diameter "d ",
fraction
e = bed porosity, fraction
Z = bed depth, cm
d = granule diameter, cm
n = overall single granule collection efficiency,
fraction
66
-------
Single granule collection efficiency includes the collec-
tion by: (1) Brownian diffusion, (2) direct interception, (3)
inertial impaction, and (4) gravity settling. The theoretical
equations using the "free surface" model for each of these four
collection mechanisms are:
4(N )
Brownian Diffusion: nn - - Sh avg (6)
NPe
= 5.04 f(er'/J Npe2/3
where nD = single granule collection efficiency, fraction
e = bed porosity, fraction
NSh = Sfterwood number, dimensionless
Np = Peclet number, dimensionless
2 - SCl-e)1 3+3(l-e)5/3- 2(l-e}2
1 - (1-e)5
2
Direct Interception: nnT = I-£• I (-8J
DI ffel
where nDI = single granule collection efficiency by direct
interception, fraction
d = particle diameter, cm
P
d = collector diameter, cm
Inertial Impaction:
c
67
-------
where TU = single granule collection efficiency by impaction,
fraction
y . = critical trajectory of the aerosol, cm
Gravity Settling: r\ = -t A (10)
where nGS = single granule collection efficiency by gravity
settling, fraction
u = terminal settling velocity of the particle,
cm/s
UG = superficial gas velocity, cm/s
A = fraction of the projected area of a single
collector particle which is available by capturing
the aerosol by settling, fraction
"A " can be taken as the minimum projected area available for
p
flow between particles. For a triangular packing it is 0.0377
and for a square packing it is 0.0871.
Model by Miyamoto and Bohn - Miyamoto and Bohn (1974) de-
rived the following equation for the particle collection in a
clean bed (diffusion only) based on the expression of a single
sphere collection efficiency by Friedlander (1957) .
Pt, = exp
6(l-e)Z NNu
dc NPeJ
where Pt, = penetration for particles with diameter "d ",
fraction
e = bed porosity, fraction
Z = bed depth, cm
d = single granule diameter, cm
NNu = N11556!1 number, dimensionless
= dckG
"V
68
-------
NPe = Peclet number, dimensionless
dc uGb
Dpe
k = mass transfer coefficient, cm/s
Dpe = effective particle diffusion coefficient in
granule layer, cm2/s
uGb = actual 8as velocity in bed, cm/s
The effective particle diffusion coefficient in granule
layer is related to particle diffusivity by:
Dpe ' Dp e * A*
where D = particle diffusivity, cm2/s
* • relative force field, dimensionless
A = tortuosity factor, dimensionless
Interstitial gas velocity can be calculated from superficial
gas velocity by:
A UG
uGb = -7- (12)
Model by Gebhart.et al. - Gebhart et al. (1973) performed
an experimental study on the filtration of aerosol particles in
the 0.1-2 ym size range by a packed bed of glass beads. Based
on their data, they proposed an empirical equation for aerosol
penetration through packed beds of spheres under conditions at
which Brownian diffusion dominates:
D 2/3 Z
Ptd = exp
-6.39
u 2/3 CO.Sd
*Gi
(13)
where Pt_, = particle penetration for particles with diameter
"d ", fraction
D = particle diffusivity, cnf/s
P
69
-------
e =
u
Gi
dc
Z
bed porosity, fraction
interstial gas velocity
collector diameter, cm
bed depth, cm
, cm/s
Balasubramanian and Meisen (1975) showed that this equation
may be derived independently from Wilson and Geankoplis1 (1966)
correlation for mass transfer coefficients. Wilson and Geankoplis
correlation is:
kG = 1.09 (uG/e) Npe~2/3 (14)
where k = mass transfer coefficient, cm/s
G
Np = Peclet number, dimensionless
. dc "G
DP
Equation (16) is valid for 0.35< e < 0.75, 0.0016 < NRe < 55,
and 950 < N~ < 70,600. The equation derived by Balasubramanian
and Meisen is:
Pt, = exp
-2.06
(l-O D 2/3
e uG2/3 (0.5dc)
(15)
This equation is more general than the equation by Gebhart et al.
After substituting e = 0.385 (bed porosity used by Gebhart et al.)
equation (17) reduces to equation (15).
Model by Bohm and Jordan - Bohm and Jordan (1976) , using
capillary flow for describing the behavior of sand bed filters,
derived an expression for particle penetration.
Pt , = exp -
/2k T ef + V « pp e dc + '
\
\3Tr d UG 36 UG
T dp2 pp UG\ 4 f'Z~
. i
18e / yG dc2
(16)
70
-------
where Ptd = particle penetration for particles with
diameter "d ", fraction
k = Boltzmann's constant
= 1.38 x 10"16 erg/°K
T = absolute temperature, °K
d = particle diameter, cm
e = bed porosity, fraction
UG = superficial gas velocity , cm/s
g = gravitational acceleration, cm/s2
d = collector diameter, cm
w
p = particle density, g/cm3
y^ = gas viscosity, g/cm-s
Z » bed depth, cm
f = d /d > dimensionless
c o
d = initial capillary diameter, cm
The first part of the exponent stands for diffusional
deposition, the second for gravity settling and the third for
inertial impaction. Interaction terms between the three col-
lection mechanisms are neglected.
Model by Goren - Goren (1977) derived a semi-empirical
equation for granular bed collection efficiency. The model
is based on collection by individual granules:
Ptd = exp - - (1-e) — n
where: Ptj = particle penetration for particles with
diameter "d ", fraction
e = bed porosity, fraction
Z = bed depth, cm
d = granular diameter, cm
n = overall single granule collection efficiency, fraction
71
-------
Goren ran a small scale experiment with 2 mm diameter gran-
ules as bed material. He then derived an expression for "n" from
data and equation (17). Collection by diffusion, settling, and
impaction were considered. The expressions are:
(18)
(19)
(20)
where n~, rir-c* HT = single granule collection efficiency due
U bo I
to diffusion, gravity settling, and impac-
tion, respectively, fraction
Np = Peclet number, dimensionless
d = granule diameter, cm
Up = gas velocity, cm/s
D = particle diffusivity, cm2/s
u = terminal settling velocity, cm/s
d = particle diameter, cm
g = gravitational acceleration, cm/s2
\ir = gas viscosity, g/cm-s
C' = Cunningham slip factor, dimensionless
P = particle density, g/cm
K = inertial parameter, dimensionless
Model by Westinghouse - Westinghouse Research Laboratories
(Ciliberti, 1977) also developed a mathematical model. They took
into account three collection mechanisms: impaction, interception,
and diffusion. Their equation is:
72
-------
Pt, = exp
Kn Z /d«\ 7 D Z
-5.8 -2— -3.75 (J£) -5. - 3.752 —E
dc VV dc UT* V
*~ C C (i C
C21)
where Ptd » penetration for particle diameter "d ", fraction
K = inertial parameter, dimensionless
dc = collector diameter, cm
A = bed depth, cm
D = particle diffusivity, cm2/s
UG = superficial gas velocity, cm/s
Model by Schmidt, et al. - Schmidt, et al. (1978), used the
semi-empirical theories of Johnstone and Roberts (1949) for
diffusion, Friedlander (1957) for interception, Jackson and Cal-
vert (1966) for inertial impaction, and Ranz (1951) for gravity
settling, and proposed the following equation for granular bed
collection efficiency.
Ptd = exp -7.5 (1-E) L- (nD + nDI + nz + nGS)l (22)
c -I
where Pt, = particle penetration for particle diameter d ,
fraction
d = collector diameter, cm
Z = bed depth, cm
e = bed porosity, dimensionless
nD » nCIt
= single granule collection efficiency due to diffu-
sion, direct interception, impaction, and gravity
settling, respectively, fraction
n
I , nGS
Single granule collection efficiencies for various collection
mechanisms were obtained from:
2 038 N
n = _8_ + Re (23)
D vt -5/8
NPe Npe/
73
-------
d l2
n = i 45 | _P_) (24)
DI 1'**\A
c
= 3.97 Kp (25)
= ^ (26)
UG
where d = granule diameter, ym
d = particle diameter, urn
K = impaction parameter, dimensionless
u. = terminal settling velocity of particles, cm/s
UG = gas velocity, cm/s
Np = Peclet number, dimensionless
ND = Reynolds number, dimensionless
Ke
Stage 2 Filtration - Stage 2 filtration has been observed
experimentally by several researchers. Billings and Wilder
(1970) summarized the studies of many investigators concerned
with the cake formation process during the initial stage of
filtration. They concluded that aerosol deposition occurs pri-
marily on previously deposited particles.
Based on this observed dendrite-like growth, Payatakes
and Tien (1976) proposed a model describing the dendrite growth
over the entire filtration period. This model was expanded
and revised by Payatakes (1977).
To express the growth process and to describe the dendrite
configuration, Payatakes (1977) divided the space adjacent to
the collector surface into layers of thickness "d " by planes
which are all parallel to a plane tangential to the collector
surface. He numbered them in ascending order; i.e., the first
layer is immediately adjacent to the collector surface (Figure 24)
The dendrite configuration is idealized with the convention
that if a particle of the dendrite structure has at least half
of its volume in the k'th layer it is assumed to lie entirely
74
-------
(a) at t = 0
(b) t = >0
in
A.
T
me =
m7 =
gm ,
m1 = i
(c) Actual dendrite
(d) Idealized dendrite
Figure 24. Dendrite initiation, growth and idealization of the dendrite
configurat ion.
-------
in the k'th layer. He also assumed that dendrite particles are
of uniform size.
By considering both the radial and angular contributions
to deposition, Payatakes (1977) set,
d m, (0) (0) (r) (r)
-Ji= 9k-l,k + 0k,k + 0k-l,k +0k,k
dt
where m, = expected particle number in the k'th layer of a
dendrite, number
0(0) = rate of increase of mfc by deposition on particles
occupying the (k-l)'st layer due to the angular
flow component, number/s
=
= rate of increase of m, by deposition on particles
already occupying the k'th layer, due to the
angular flow component, number/s
frl
®v i v = rate of increase of m, by deposition on particles
K .L ) K K
occupying the (k-l)'st layer due to radial flow
component, number/s
£>y L = rate of increase of m, by deposition on particles
K , K K
already occupying the k'th layer due to the radial
flow component, number/s
In the formulation, terms involving addition of particles to
the kth layer by deposition onto particles occupying the (k+l)'st
layer are neglected. The solution to equation (27) is»
mk = f Cki (9) exp [a bi (9) tJ' for k = 1, 2, 3, ... (28)
where mk = expected particle number in the kth layer of a
dendrite
t = time measured from the instant of deposition of
the first particle of the dendrite, seconds
a = rate of particles approaching a clean fiber per
unit length = d ur. n
C Lil
76
-------
0 • angular cylindrical coordinate, measured counter-
clockwise from the down stream stagnation point
d » collector diameter, cm
uGi = interstitial gas velocity, cm/s
n = particle number concentration, number/cm
"Cv- (Q)" and "b. (0)" are functions defined by the following
& -L J-
equations:
C = 1 (29)
11
c = __!^ (30)
2 1 fK, -K, 1
C = ——
b,- b,
n aj (32)
Jlk b,-b3
^i n ^— , for k = 3,4..., (33)
2
-------
where p = maximum number of particles in the (k+l)'st layer
which can be attached directly to the same particle
of the k'th layer
4>?s-l = function defined so that "a 4>>S]M is the rate of
i > 3 1»J
increase of "m." by deposition on a particle in the
i'th layer due to the flow component in the "s"
direction (s = r = radial direction, s =0= angular
direction).
a _ , (6) + , (r)
ak ~ * k-l,k + * k-l,k
Equation (28) coupled with the assumed flow field (e.g.;
Happel's free surface model, Kuwabara flow field, etc.) can be
used to predict the increases in filtration efficiency and pres-
sure drop for the filter. Payatakes (1976 a,b) presented some
sample calculations by applying the model to pure interception.
Payatakes and lien's model described the rate of dendrite
growth. It did not explain the reason for dendrite formation.
Wang et al. (1977) and Tien et al. (1977) proposed a concept,
the shadow effect, for dendrite growth.
They hypothesized that once a particle is deposited, it
creates a shadow area around itself on the collector surface,
within which no subsequent particle deposition may take place.
This is represented by arc B'BB" for the deposited particle "A"
shown in Figure 25.
The creation of shadow areas by deposited particles has two
consequences. Since there will be no deposition with any shadow
area, it means that particle collection takes place at a discrete
position along a collection surface, the deposited dust cannot
be in the form of a smooth coating.
The second consequence arises from the fact that with the
creation of shadow area, subsequent approaching particles which
would have deposited within the shadow area had there been no
deposition, now attach themselves to the deposited particle. This
results in the formation and growth of chain-like particle
dendrites.
78
-------
Limiting Trajectory
Angle Covering shadow area
to
Figure 25. Particle deposition on single collector
-------
The magnitude of the shadow area is a function of many
variables including the location of its deposition, particle
diameter, and the inertial parameter. They performed a simula-
tion by using this concept and the trajectories of random par-
ticles to verify this concept.
Gaseous Pollutants
It is feasible to use various bed materials to remove metal
vapors and other gaseous pollutants. The mechanisms expected
to be important in removing gaseous pollutants with a granular
bed include:
1. Adsorption
2. Absorption
3. Chemical reaction
4. Condensation without nucleation
Adsorption -
The Chemical Engineers' Handbook gives a detailed treatment
of this subject. The following is a brief discussion. Adsorption
involves the interphase accumulation or concentration of sub-
stances at a surface or interface. The process can occur at an
interface between any two phases, such as gas-solid, gas-liquid,
liquid-solid, or liquid-liquid interfaces. The material being
concentrated or adsorbed is the adsorbate, the the adsorbing
phase is termed the adsorbent.
There are three principal types of adsorption: electrical
attraction of the solute to the adsorbent, van der Waals attrac-
tion and chemical adsorption. Adsorption of the first type falls
within the realm of ion-exchange and is often referred to as ex-
change adsorption. Van der Waals attraction is generally termed
"physical" adsorption, a term which has come to represent cases
in which the adsorbed molecule is not affixed to a specific
site at the surface but is free to undergo translational movement
within the interface. Adsorption of this type is sometimes
referred to also as "ideal" adsorption. If the adsorbate under-
goes chemical interaction with the adsorbent, the phenomenon is
referred to as "chemical" adsorption, "activated" adsorption, or
80
-------
"chemisorption." Chemically adsorbed molecules are considered
not to be free to move on the surface or within the interface.
Performance of a solid sorbent depends upon four factors.
1. Stoichiometric capacity of the solid. This is the ul-
timate capacity of the sorbent for the sorbate, which may or may
not be fully utilized under actual process conditions.
2. The phase equilibrium, which influences the efficiency
with which that capacity is reached, and in many cases controls
the actual capacity of the surbent.
3. The rate behavior including the mechanism and resistance
controlling the mass transfer rate.
4. The process arrangement and its effect on the material
balance.
Equilibrium Behavior - When the concentration of the sorbate
remaining in the gas phase is in a dynamic equilibrium with that
at the surface of the solid, there is a definite distribution of
the sorbate between the gas and solid phases. One form used to
depict this distribution is to express the quantity "q " as a
6
function of "c" at a fixed temperature. The quantity "q " is
C
the amount of sorbate adsorbed per unit weight of sorbent and
"c" is the concentration of sorbate remaining in gas phase at
equilibrium. An expression of this type is termed an adsorption
isotherm. The adsorption isotherm is a functional expression
for the variation of adsorption with adsorbate concentration
in the gas at constant temperature. Usually the amount of ad-
sorbed material per unit weight of adsorbent increases with
increasing concentration, but not in direct proportion (Figure 26),
Several types of isothermal adsorption relations may occur.
The most common relationship between "q " and "c" is obtained
for systems in which adsorption from the gas leads to the depo-
sition of an apparent single layer of sorbate molecules on the
surface of the solid. Occasionally, multimolecular layers of
sorbate may be adsorbed. The single layer adsorption can be
described by the Langmuir adsorption model and the multilayer
adsorption by the Brunauer, Emmett, Teller (BET) model.
81
-------
a. Langmuir
c
BET
Figure 26
Typical isotherms for
Langmuir and BET adsorption
patterns.
82
-------
The Langmuir treatment is based on the assumption that
maximum adsorption corresponds to a saturated monolayer of sor
bate molecules on the adsorbent surface, that the energy of ad-
sorption is constant, and that there is no transmigration of
adsorbate in the plane of the surface. Figure 26 is a typical
isotherm for the Langmuir pattern. The Langmuir isotherm is:
qe - - (36)
e 1 + Kc
where q = the amount of sorbate adsorbed per unit weight of
C
adsorbent in equilibrium with concentration "c"> mol/g
Q° = number of moles of adsorbate adsorbed per unit
weight of adsorbent in forming a complete monolayer
on the surface > mol/g
K = equilibrium constant> £/mol
c = concentration or partial pressure of adsorbate
in gas phase, mol/fc or mm Hg
The BET model assumes that a number of layers of adsorbate
molecules form at the surface and that the Langmuir equation
applies to each layer. A further assumption of the BET model
is that a given layer need not complete formation prior to the
initiation of subsequent layers. For adsorption from the gas
phase with the additional assumption that layers beyond the
first have equal energies of adsorption, the BET equation takes
the simplified form;
(c -c) [1 + (B-l)(c/c )]
s =
83
-------
where c = saturation concentration of the adsorbate, mol/g
c = measured concentration in gas at equilibrium, mol/4
B = a constant expressive of the energy of interaction
with the surface, 2,/mol
Q° = number of moles of adsorbate adsorbed per unit
weight of adsorbent in forming a complete monolayer
on the surface, mol/g
q = number of moles of adsorbate adsorbed per unit
weight at concentration "c", mol/g
One other question for isothermal adsorption, the Freundlich
or van Bemmelen equation, has been widely used for many years.
This equation is a special case for heterogeneous surface ener-
gies in which the energy term, "b", in the Langmuir equation
varies as a function of surface coverage, "q ", strictly due to
"
variations in heat of adsorption. The Freundlich equation has
the general form:
= KF c
1/n
where "Kp" and "n" are constants and n > 1.
Rate Processes - There are essentially three consecutive
steps in the adsorption of materials from gas by porous adsorbents,
any one of which may be a rate determining step in a certain re-
gion of operating conditions.
The first step is the transport of the adsorbate through
a surface film to the exterior of the adsorbent. The transport
rate for adsorbate between the bulk of the gas phase and the
outer surfaces of the sorbent granules is governed by the mole-
cular diffusivity and, in turbulent flow, by the eddy diffusivity
which controls the effective thickness of the boundary layer.
84
-------
One may assume that the concentration of sorbate at the point
in the hydrodynamic boundary layer immediately adjacent to the
external surface of the particle is in equilibrium with the average
solid-phase concentration on the internal surfaces. This condition
may be stated algebraically as:
- kG a £- (c - ce) (39)
where kg « mass-transfer coefficient, cm2/s
a = effective area for mass transfer across the fluid
film per unit volume of bed, cm2/cm3
e « void fraction, fraction
p = density of the solid in the bed, g/cm3
c « concentration of the sorbate in bulk gas phase,
g/cm3
c = concentration of sorbate in bulk gas phase in
equilibrium with the coexisting solid phase
concentration, g/cm3
q = concentration of sorbate on solid surface, g/cm3
For packed bed, Wilke and Hougen (1945) gave the following
equation for evaluating the mass-transfer coefficient:
0.51 0.16
10.9 ur U-E) / Dr \/Dr Pp\
k a = * ( —M-5—-) (4°)
dc ^dc UG/V »G '
where UG = superficial gas velocity, cm/s
e = void fraction , fraction
d = granule diameter, cm
DG = gas phase diffusivity, cm2/s
PG = gas phase density, g/cm3
= gas phase viscosity, g/cm-s
85
-------
The second of the three consecutive steps in sorption by
porous adsorbents, with the exception of a small amount of
adsorption that occurs on the exterior surface of the adsorbate
after transport across the exterior film, is the diffusion of
the adsorbate within the pores of the adsorbent. The driving
force approximation for pore diffusion is:
- q) (41)
where k = pore diffusion coefficient, cm /s
a = outer surface area of particles per unit bed
volume, cm2 /cm
5 = intraparticle void ratio, dimensionless
q = concentration of sorbate on solid surface, g/cm3
q = local concentration of sorbate in the solid phase
G
that prevails at the outer surface and is assumed
to be in equilibrium with the coexisting gas phase
at concentration "c", g/cm3
According to Vermenlen and Quilici (1970) :
60
<42)
where D ____ = pore diffusity, cm2/s
po i c
d = solid diameter, cm
e = bed void fraction, dimensionless
Pore diffusivity can be expressed as:
= i F-L- /-^-IL\ + L.V
x [4? \2 RT; DJ
86
D
-------
where D _ « pore diffusivity, cm2/s
pore
5 = internal void fraction of the solid, fraction
r = average pore radius, cm
X = tortuosity factor (usually between 2 and 6),
dimensionless
R = universal gas constant, J/gmol-°K
T = absolute temperature, °K
M - molar weight, g
Dp = gas phase diffusivity, cm2/s
The third and final step is the adsorption of the adsorbate
on the interior surfaces of the adsorbent (e.g. porous granules)
The following reaction schematic for the adsorption process is
considered:
(sorbate) + (sorbent) -*• (sorbate • sorbent)
If a Langmuir-type adsorption equilibrium is assumed, a
general expression for the rate of adsorption at the solid
surfaces may be given as:
d (sorbate-sorbent) _
(sorbate) (sorbent) -
dt
(sorbate.sorbent)
(44)
where K = Langmuir equilibrium constant, A/mol
k - rate constant for second-order surface reaction
controlled kinetics, A/mol-s
If the ultimate monolayer capacity of the adsorbent for the
adsorbate is designated by the term Q°, then (Q°-q) represents
the unused capacity of the adsorbent. Substitution into
equation (44") yields:
If - k [c CQ°-q) - |]
87
-------
The gas-solid heterogeneous reaction may be either a non-
catalytic or a catalytic reaction. In this report, we will con-
sider the heterogeneous, noncatalytic reaction. The treatment
of this type of reaction requires the consideration of mass trans-
fer between phases, the contacting patterns of the reacting phase,
and the reaction kinetics. Levenspiel (1972) and the Chemical
Engineers' Handbook present a detailed treatment of this sub-
ject. The following is a brief summary of those discussions.
A heterogeneous non-catalytic reaction may be represented
by:
A (gas) + b B (solid) •* Products
The products may be gas, solid or both. There are several models
available to describe the progress of the above reaction. The
unreacted core model seems to work reasonably well in a wide
variety of situations. This model considers that the reaction
occurs first at the outer skin of the solid. It can leave behind
both completely reacted material, and inert solid material residue.
Thus, at any time there exists an unreacted core of material which
shrinks in size during reaction. Two different cases may be
considered for this model. The first assumes that the continuous
formation of solid product (inert residue) without its flaking
off would maintain the particle size unchanged. In the second
case the particle size changes as the reaction progresses owing
to the formation of gaseous products, flaking off of the solids,
etc.
Unreacted Core Model for Spherical Particles of Unchanging Size
This model visualizes the reaction occurring in three
successive steps. Either one of the three steps may be rate
determining:
Step 1; Diffusion of gaseous reactant "A" through the film sur-
rounding the solid to the exterior surface of the solid.
Step 2: Penetration and diffusion of the reactant "A" through
the blanket of residue to the surface of the unreacted core.
-------
Step 3: Reaction of gaseous reactant with solid at the reaction
surface.
If the resistance of the gas film is the controlling factor,
the reaction rate is equal to the diffusion rate of the gas reac-
tant from the bulk phase to the interface. The chemical reaction
can be assumed to be instantaneous. Thus, the concentration of
the gas at the solid surface is zero and the concentration driving
force for diffusion is constant at all times during the reaction
of the particle. In terms of the shrinking radius of the unreac-
ted core, the reaction rate is:
- b kr c (45)
dt G
where p = density of reactant "B", g/cm3
o
r = radius of unreacted core, cm
R « original radius of the reacting particle, cm
b = stoichiometry constant, dimensionless
kg = mass transfer coefficient, cm2/s
c = concentration of gas reactant "A" in the bulk gas
phase, g/cm3
t = time, s
Rearranging and integrating, we find how the unreacted
core shrinks with time,
t = Ps R I i - i l£l 1 (-46)
3b kr. c
Let the time for complete reaction of a particle be "T".
Then by taking "r =0" in equation (46), we find,
P
T =
3b k
* (47)
89
-------
The radius of the unreacted core expressed in terms of fractional
time for complete conversion is obtained by combining equations:
(48)
^R
where XD = fraction of reactant "B" converted into product
D
By using the same approach, the integrated rate equations
under other conditions are:
Diffusion through residue layer c< itrols:
^ = 1-3 (l-xB)2/3 + 2 (l-xB) (49)
PS R2
T = 5 (50)
6 b D c
e
where D = effective diffusion coefficient of gaseous reactant
in the residue layer, cm2/s
Chemical reaction controls:
t 13
- = 1- d-xR) (51)
T B
P R
(52)
b k
where k = first order rate constant for the surface reaction.
Unreacted Core Model for Shrinking Spherical Particles
When no adherent residue forms, the reacting solid particle
shrinks during reaction, finally disappearing. For a reaction
of this kind we visualize the following three steps occurring in
succession.
Step 1: Diffusion of reactant "A" from the main body of gas
through the gas film to the surface of the solid.
90
-------
Step 2: Reaction on the surface between reactant "A" and solid.
Step 3: Diffusion of reaction products from the surface of the
solid through the gas film back into the main body of gas.
As with constant size particles, the following rate ex-
pressions result when one or the other of the resistances control
Diffusion through gas film controls:
Small particle (Stokes regime)
£ = 1- d-xB)2/3 (53)
T
T .
2 b DG c
where y. = mole fraction of reactant "A" in gas phase
DG = gas diffusivity, cm2/s
Large particle (UG = constant)
£ = 1- (l-xB)°<5
a/a
T = (constant) - —
Chemical reaction controls: When chemical reaction controls,
the behavior is identical to that of particles of unchanging size;
therefore, equations (51) and (52) will represent the time-conver-
sion behavior of single particles, both shrinking and of constant
size. With this information on reaction kinetics, we can determine
the required granular bed size for various gas-solid contacting
schemes (fixed bed batch process, continuous moving bed, etc.).
91
-------
Cake Filtration
Granular bed filters usually have larger pore sizes and greater
thickness than the fibrous filter. Whether internal cakes will
form depends largely on the granule size. If the granules of the
bed are large, dendrites will not bridge to create an internal
cake. On the other hand, if the grains are sufficiently small,
dendrites will bridge to form an internal cake.
Lee (1975) called the internal cake a rooting cake and it
is the foundation to support the formation of a surface cake.
Once a surface cake is formed, filtration efficiency no longer
depends upon the depth of the granular bed but rather on the
thickness and structure of the surface cake. The cake filtration
results in a much higher efficiency than the original granular
bed. This is because particle collection by sieving becomes
a more important collection mechanism.
Leith, et al. (1976) and Leith and First (1977) studied
high velocity cake filtration of fabric filters. Three mechanisms
were described by them by which particles can pass through a
fabric filter or a granular bed. The three mechanisms are:
1. Straight through penetration
2. Seepage or bleeding penetration
3. Pinhole plug penetration
In straight through penetration, particles pass through the
filter without stopping; i.e., they are not collected by the
filter. Once a particle lands on or in the filter, it needs not
necessarily remain at its point of initial impact. As the dust
deposit builds up, the dust may work its way through from the
dirty to the clean side of the filter. This may result from the
drag force exerted on the particle deposits by the gas moving
past. Penetration of this sort is called seepage or bleeding.
The pinhole plug mechanism postulates a plug of deposited par-
ticles dislodged from the dust deposit and moves out of it,
leaving behind a pinhole. Figure 27 is a schematic representa-
tion of penetration mechanisms.
The size distribution of particles passing through the fil-
ter by the straight through mechanism should reflect a dependence
92
-------
I
I
\
I
6
i
t
O
I
I
i
I
O
ii
Straight
Through
Seepage
Pinhole
Plugs
Figure 27. Schematic representation of fly ash emission
mechanisms.
93
-------
on the forces causing particles to be collected there: inertia,
interception, diffusion, gravity, etc. However, the size distri-
bution of the particles which pass through by seepage or pin-
hole plugs should be the same as the size distribution of the
deposited dust; that is, very close to the size distribution of
dust fed to the filter.
Using a series of tagged dusts, the proportion of total
dust emitted which is accountable to each emission mechanism
was measured by Leith et al . (1976) in relation to face velocity and
deposit thickness. Deposit thickness is defined as:
W C ,t
= P1
pp (1-e) Pp -e
where X = dust deposit thickness, um or cm
p = particle density, g/cm3
W = particulate load, g/cm2
C . = inlet particle concentration, g/cm3
Up = superficial gas velocity, cm/ s
t = time since last cleaning, s
They found significant trends in the dust penetration mech-
anism data. The time-mechanism interaction was highly significant.
As time increases and the dust deposit thickens, the mechanisms
by which dust penetrates the filter change. Straight through
penetration rapidly diminishes in importance although it is
important immediately after a filter cleaning cycle. The emitted
dust accountable to the seepage mechanisms is relatively constant
during the entire filtration cycle. The pinhole plug mechanism
rapidly rises in importance after cleaning, passes through a maxi-
mum, and then declines as the dust deposit becomes thicker and
pressure drop through the deposit increases. Figure 28 is an
illustration of the trends.
The velocity mechanism interaction was not significant. At
any fixed time, the fraction of dust emitted by each penetration
mechanism is fairly constant at all velocities tested.
94
-------
1.0
O
•H
+J
U
o
CO
CO
CO
^5
Q
H
u
0.8
o 0.6
0.4
0.2
Pinhole
Plugs
Straight
Through
I
20 40
DEPOSIT THICKNESS, ym
60
Figure 28. Fraction of total fly ash emitted by various
mechanisms as a function of deposit thickness.
95
-------
Miyamoto and Bohn (1975) studied the effect of particle
loading on granular bed filter collection efficiency. Their
results were presented in Figures 13, 14 and 15 in an earlier sec-
tion of this report.
Under contract with EPRI, Squires and his co-workers at City
College of New York have experimentally determined the collection
efficiency of a granular bed with filter cake. A final report
has recently been published (Lee et al., 1977).
Pressure Drop
Clean Bed -
Several investigators have shown that the flow through packed
beds can be described by: * 7 „ n ^
I L Uf-i L -1- ~ E J Pr1
-AP = — (58)
dc *'
where AP = pressure drop across packed bed, g/cm-s
f = friction factor, dimensionless
UG = superficial gas velocity, cm/s
PG = gas density, g/cm3
e = bed porosity, fraction
d = bed particle diameter, cm
Figure 29 is a plot of friction factor versus Reynolds num-
ber for a fixed bed. Ergun (1952) has defined a Reynolds number
d u,, p.-,
NRe ' f-^f <59>
For laminar flow with NR < 1.0; and by analogy to flow in
many other systems, we can approximate "f" by a constant divided
by "NR ." An analysis of experimental data indicates that the
constant is 150. Therefore, for laminar flow we have:
f = P (60)
This is referred to as the Kozeny-Carman equation. For a
given bed and fluid, it predicts that the flow rate is proportional
to the pressure drop, which is D'Arcys law.
96
-------
<
I
a;
o
H
U
H
U
I I I I I Illl I I
IMI
BURKE PLUMMER
10
20
30
40
N
Re
1-e
Figure 29. A comprehensive plot of pressure drop
in fixed beds.
97
-------
For completely turbulent flow, it is reasoned that "£"
should approach a constant value and that all packed beds have
the same relative roughness. The constant is found by experiment
to be 1.75, so we have:
(-AP) d e
£ = 1.75 = - - - (61)
Z UG* PG (1-0
This is called the Burke-Plummer equation.
A consideration of flow at intermediate Reynolds numbers
led Ergun (1952) to propose as a general equation:
£ = HO + 1>?5 (62)
NRe
Filter Cake Resistance to Gas Flow -
Miyamoto and Bohn (1975) studied the effect of particulate
load on pressure drop. Figures (30) through (32) show their results.
The pressure drop remained constant until the particulate load
of the filters reached a threshold load, then increased rapidly
as shown in the figures.
In granular bed filters, the flow resistance should change
little so long as the large pores are open, but will increase
when the large pores are closed by surface cake. Depending on
whether the compaction effect of the filter cake is present,
the increase of pressure drop is at a different rate.
No compaction effect - The pressure drop across the surface
cake is given by D'Arcy's law, i.e.,
G (63)
Kd
98
-------
u
3
SJ
2
O
h- 1
H
U
WH
I I ill I I I I I llll
rm i r
i i i t ii ni
BURKE § PLUMMER
i i ii ni i i i 111 in i i
10
20
30 40
N
Re
1-e
Figure 29. A comprehensive plot of pressure drop
in fixed beds.
97
-------
For completely turbulent flow, it is reasoned that "£"
should approach a constant value and that all packed beds have
the same relative roughness. The constant is found by experiment
to be 1.75, so we have:
f = 1.75 =
This is called the Burke -Plummer equation.
A consideration of flow at intermediate Reynolds numbers
led Ergun (1952) to propose as a general equation:
£ . 15° + 1.75
NRe
Filter Cake Resistance to Gas Flow -
Miyamoto and Bohn (1975) studied the effect of particulate
load on pressure drop. Figures (30) through (32) show their results.
The pressure drop remained constant until the particulate load
of the filters reached a threshold load, then increased rapidly
as shown in the figures.
In granular bed filters, the flow resistance should change
little so long as the large pores are open, but will increase
when the large pores are closed by surface cake. Depending on
whether the compaction effect of the filter cake is present,
the increase of pressure drop is at a different rate.
No compaction effect - The pressure drop across the surface
cake is given by D'Arcy's law, i.e.,
- 4P . GG (63)
Kd
98
-------
ft.
<
g
Cu
<
80
60
40
20
0-1 0.2
PARTICULATE LOAD, kg/mz
Figure 30. Effect of mean granule
diameter on pressure drop
(Miyamoto and Bohn's data),
80
60
40
20
Z-10 cm
30
» 2 cm/s
« 3 mm
70 90
0 0.1 0.2
PARTICULATE LOAD, kg/m2
Figure 31. Effect of bed depth on
pressure drop (Miyamoto and
Bohn's data).
80
60 -
40 -
20 -
0 0.1
PARTICULATE LOAD, kg/ma
Figure 32. Effect of superficial gas
velocity on pressure drop
(Miyamoto and Bohn's data),
99
-------
where Ap = pressure drop across packed bed, g/cm-.s2
yG = gas viscosity, g/cm-s
u,, = superficial gas velocity, cm/s
Z = bed depth, cm
K, = D'Arcy permeability, cm2
For laminar flow, it becomes
150 Z u^ (1-e) pc
where e = bed porosity, fraction
PG = gas density, g/cm3
d = granule diameter, cm
NR = Reynolds number , dimensionless
This is referred to as the Kozeny-Carman equation.
By combining equations (63) and (64) we obtain
150 y Z u (1-e)2
C65)
Substituting equation (57XX = Z) into equation (65) gives:
150 y C -t u2 (l-e)
C66)
where Ap = pressure drop across filter cake, g/cm-s2
V~ = gas viscosity, g/cm-s
t = time since last cleaning, s
UG = superficial gas velocity, cm/s
e = filter cake porosity, fraction
d = granule diameter, cm
\*
p = particle density, g/cm3
C • = inlet particulate concentration, g/cm3
100
-------
Equation (66) predicts that at constant inlet conditions, the
pressure drop across the surface cake varies linearly with time
as long as the porosity of the filter cake remains constant (no
compaction effect present). It also predicts a linear relation-
ship between pressure drop and the square of the superficial
gas velocity, after the same time period, for equal loadings and
bulk density of the filter cake.
Compaction Effects - Leith et al (1976) also have studied the
compaction of filter cake and its implications. They found that
Kozeny-Carman's equation yields excellent results for powders
which are compressed to a specified porosity provided the particles
are isometric.
Compaction of the cake is a complicated phenomenon which is
related to the increased drag exerted by the gas on the cake.
Although the dust on the surface of the cake is under negligible
mechanical stress due to this drag, the dust below the surface
must support the drag experienced by the dust layer above it. At
the bottom of the cake, the compressive stress is the greatest
and equals the total pressure drop across the cake, plus the stress
due to the impact of new particles.
Limited studies have been done to determine the effects of
increased velocities on permeability. Stephan et al. (1960)
observed a 60% change in this factor when velocity was increased
from 1.8 to 3.0 m/min with clean air, and the change was noted
to take place in 5 discrete collapses. The cake was initially
deposited at 1 m/min (3.4 ft/min). Borguardi et al (1968)
reported a similar change with the same dust (fly ash). These
studies provide some insight into cake collapse. However, if
the cake is formed at high velocity, and the velocity remains
constant, the compaction of the cake may proceed more gradually
due to tighter initial packing.
Orr (1966) indicated that compaction of cake can take place
by three mechanisms:
1. sliding,
2. elastic and plastic deformation,
and 3. fragmentation.
101
-------
The latter is unlikely except for extremely fragile particles
due to the relatively small compressive stresses. The most likely
mechanism is sliding which is opposed by frictional and cohesive
forces between particles. Compaction effects are often
described by the following exponential expression:
e = ei exp (-a F) (67)
where e = final void fraction, fraction
e. = initial void fraction, fraction
a = constant, cm-s2/g
F = compression stress, g/cm-s2
Differentiating equation (67) gives:
3F - - a e (68)
Differentiating equation (57) yields:
dx - + *_ + e (69)
P P (
D'Arcy's equation is
dp
Combining equations (69) and (70) we obtain:
K dW W de
dp =
From the Kozeny-Carman equation
K =
d 150
(70)
102
-------
Combining equations (71) and (72) yields
dp = -
150 »G U
dc2 pp
For a small increment of cake
de
(73)
dF = - dP
Combining equations
dW _ k e2 + W
HT ~ (e-1)
where k = constant, g/cm2
d 2 p
c "p
150 a y~ UG
Solving equation (75) with boundary condition]
e = ei at W « 0, yields
W
= e2-2e-2(l-e)£n(l-e)-(l-e)
2ei-ei2
(74)
(75)
(76)
Since the total compressive stress on the cake is equal to "p",
equation (67) can be rewritten as,
e - e^^ exp [ -a (Ap)]
Substituting equation (76) into equation (77) yields,
(77)
J
. 2
[
(78)
103
-------
Since W = C .ur t
pi G
We have
-2e. e -2
in
1-e.
1 (80)
A plot of pressure drop versus time for cake buildup can be a
curve fitted to equation (80) and the constants a, e^, and k
evaluated by regression analysis.
104
-------
TABLE 14. AVAILABLE EQUATIONS FOR THE PREDICTION OF
PARTICLE COLLECTION IN A GRANULAR BED
Investigator
Equation
Notes
Jackson and Calvert
(1968)
Pt, = exp
Impact ion only.
Paretsky et al.
(1971)
Ptd = exp
2 e
Miyamoto and Bohn
(1974)
exp
Z NNu
dc
Collection by
diffusion only
Gebhart et al.
(1973)
Ptd = exp
D 2/8
-6.39 -2 r
Collection by
diffusion only
Bbhm and Jordan
(1976)
Pt, = exp
d
d
/2 k T e f + ^> 8
\37T d U
e d
36
. T V PP UG
18e
-)
4 £' Z
Collection by
diffusion.
Gravity settling
and impact ion.
Continued
-------
TABLE 14. (continued)
Investigator
Equation
Notes
Goren
(1977)
Pt = exp
- (1-e) -
O.75
l250 K
2'25
Collection by impaction,
gravity settling, and
diffusion
o
00
Westinghouse
(Ciliberti, 1977)
Pt = exp
f
- 5.
K Z
- 3.75
3.572
Collection by impaction,
interception, and diffusion
Schmidt, et al .
(1978)
Pt,, =
7.5 Z
* 2.038
(d \
-£)
c/
Collection by diffusion,
interception, impaction,
and gravity settling
+ 3.97 K +
u 1
-^
U-I
-------
and» Ptd "
where n = single granule collection efficiency, fraction
Pt^ = particle penetration, fraction
and subscripts D, DI, I and GS refer to diffusion, direct inter-
ception, and gravity settling, respectively
As can be seen from Figure 33 for this particular granular
bed filter, their equation predicts that the collection efficiency
will be very low. This is not in agreement with McCain's data.
This discrepancy may result from Paretsky et al . basing their
equation on collection by an isolated sphere. At a gas velocity
of 80 cm/s (superficial gas velocity in the granular bed during Mc-
Cain's tests), the dominant collection mechanism is inertial im-
paction. For particle collection by inertial impaction onto a
single spherical collector, it may be generally assumed that there
will be no collection if the inertial impaction parameter is below
a critical value. The critical value of impaction parameter is
about 0.083 for an isolated spherical collector. For the test
conditions in McCain's tests, this is equivalent to the impaction
parameter of a 10 umA diameter particle. Therefore, for particles
with diameters smaller than 10 ymA, there should be no collection
by impaction.
Bohm and Jordan (1976) derived their equation by visualizing
the granular bed as a system of parallel capillaries. The ratio
of granule diameter to initial capillary diameter, "f", changes
during the filtration period due to accumulation of collected
particles. According to Bohm and Jordan,
6.5 < f < 10
109
-------
The predicted particle penetration for f = 6.5 and f1 = 10
is plotted in Figure 33. By assuming there was no surface cake
and the pressure drop across the bed was 80% of the overall pres-
sure drop, the bed porosity was estimated to be 0.25 (by Ergun's
equation). With this value for bed porosity, Bohm and Jordan's
equation predicted too low a particle penetration. If we assume
f = 0.75 the prediction will match the data. However, this
assumption is unrealistic because for f1 < 1, the capillary dia-
meter will be greater than the granule diameter.
The predicted penetration based on Goren's model for e = 0.25
is higher than measured.
As mentioned earlier, the dominant collection mechanism was
inertial impaction for the operating conditions of the filter
during McCain's tests. Both Miyamoto and Bohm's equation and
Gebhart et al.'s equation are for particle collection by diffu-
sion only. Therefore, these two equations are not suitable for
comparing with McCain's data.
For collection by impaction, Westinghouse's model and Schmidt
et al.'s model reduce to:
7
Westinghouse: Pt, = exp (-5.8 — K )
a d p
Schmidt et al.: Pt, = exp [-29.85 (1-e) — K ]
a d p
c
Except for the constant, these equations are identical to
that of Jackson and Calvert (1968) . Predictions by these equa-
tions are compared with McCain's data in Figure 33 . Bed porosity
is assumed to be 0.25. As can be seen, Westinghouse's model
slightly overestimated the penetration and Schmidt et al.'s
model underestimated particle penetration.
DATA REPORTED BY HOOD
Hood (1976) reported the evaluation of the Combustion Power
Company's moving gravel bed filter on the control of particulate
emissions from a hog-fuel fired boiler. The gravel bed filter
110
-------
was a prototype unit with suggested capacity of 1,133 Am3/min
(40,000 ACFM). The bed was packed with an intermediate size
gravel which was retained on a 3.2 mm (1/8 in.) wire mesh and
passed a 6.4 mm (1/4 in.) mesh screen. The bed was a single down-
flowing annulus 2.6 m (8.5 ft) O.D. and 1.8m(6 ft) I.D.
During sampling the unit was operated at a flow rate of
1,558 m3/min (55,000 ACFM). The gas temperature was 177°C (350°F)
Particle size distribution and concentration were sampled
with cascade impactors. Particle penetration was calculated from
the cascade impactor data.
Figure 34 shows the comparison between Hood's results and
predictions by available design equations. The bed porosity was
calculated to be 0.25 (from Ergun's equation). As can be seen
from Figure 34, none of the available design equations agrees
with the measured performance.
Under the sponsorship of ERDA, Combustion Power Company
conducted experimental studies on their GBF system to correlate
the collection efficiency of the GBF with mechanical and process
parameters. Parameters studied included superficial gas velocity,
dust loading, particle size and distribution, granule diameter,
granule circulation rate, and bed thickness.
The GBF was a pilot unit and was operated at ambient tempera-
ture. Redispersed hydrated alumina was used as test dust.
A.P.T. has acquired some data with cascade impactors on this
GBF system. Figures 35 through 37 show part of the data along
with predictions by available design equations. The model pre-
dictions do not agree with the data.
DATA BY KNETTIG AND BEECKMANS
As mentioned earlier, all of the available equations are
for the prediction of particle collection by clean beds. The
granular bed filter data reported by McCain and by Hood are data
obtained on industrial installations. The beds will not be clean.
In the following sections, data obtained on laboratory scale clean
granular bed filters will be used to test the design equations.
Ill
-------
1.0
0.5
0.1
0.05
0.01
HOOD'
DATA
A. JACKSON & CALVERT
B. PARETSKY, ET AL.
C. BOHM & JORDAN, f'«10
BO'HN ft JORDAN, f "6.5
E. GOREN, e= 0.25
F. WESTIN6HOUSE
G. SCHMIDT, e= 0.25
0.1
0.5 1.0 5
AERODYNAMIC PARTICLE DIAMETER, umA
10
Figure 34. Comparison of Hood's data with predictions by available
design equations.
1.0
0.5
o
3
u.
. 0.1
S
£ 0.05
0.01
GOREN, E= 0.35
WESTINGHOUSE
UQ = 81 cm/s
d » 0.2 cm
Z = 20.2 cm
GRANULE CIRCULATION
RATE = 0.8 kg/kg air
\
SCHMIDT, ET AL. \
e= 0.35
\
\
DATA
0.1
1 i I
10
0.5 1.0 5
AERODYNAMIC PARTICLE DIAMETER, umA
Figure 35. Experimental and predicted performance of CPC GBF (A.P.T. data).
112
-------
1.0
0.5
"• o.i
i '
1
£ 0.05
0.01
WESTINGHOUSE
SCHMIDT, ET AL
e = 0.35
GOREN, e= 0.35
GRANULE CIRCULATION
X RATE =
"X 1.6 kg/kg
xx air
N.
A.P.T. DATA
u. = 41 cm
d = 0.2 cm
Z - 20.2 cm
\
GRANULE CIRCULATION \
RATE = 0.8 kg/kg air
0.1 0.5 1.0 5 10
AERODYNAMIC PARTICLE DIAMETER, umA
Figure 36. Experimental and predicted performance of CPC GBF (A.P.T. data).
1.0
0.5
o
5
S
o.i
0.05
0.01
GOREN, e= 0.35
WESTINGHOUSE
SCHMIDT, ET AL.
e = 0.35
u. = 41 cm/s
d = 0.2 cm
Z = 40.6 cm
0.1
10
0.5 1.0 5
AERODYNAMIC PARTICLE DIAMETER, umA
Figure 37. Experimental and predicted performance of CPC GBF (A.P.T. data).
113
-------
Knettig and Beeckmans (1974) studied the capture of mono-
disperse aerosol particles in the size range of 0.8-2.9 urn in
a screen supported and in a grid supported fixed bed of 425 urn
glass beads. Test results showed a linear relationship between
collection efficiency, expressed in transfer units, and bed
height. Impaction appears to have been the primary collection
mechanism because collection efficiency increased with both
superficial gas velocity and aerosol particle size. Transfer
units are related to penetration by:
NTU = - In Ptd (81)
In terms of number of transfer units, various design equa-
tions become:
Jackson and Calvert: ^- = - 2. (82)
Z d
Paretsky et al . : » (i-e) IL_ (83)
Z 2 d_
Goren: = 1 (i-e) Jl (84)
Z 2 dc
Bohm and Jordan:
2 TT f • Kp (85)
9 e d
Westinghouse: =2.9-2- (86)
Z d
c
Schmidt, et al . : — = 29.85 (1-e) -2- (87)
Table 15 compares data with predictions. None of the pre-
dictions agree with the data. Jackson and Calvert's equation
(1965) and Bohm and Jordan's (1976) equations overpredict effi-
ciency. Paretsky et al.'s equation predicts no particle
114
-------
TABLE 15. COMPARISON OF KNETTIG AND BEECKMAN'S DATA AND PREDICTIONS
Support Superficial
Gas Velocity
cm/s
Screen
Support
Grid
Support
8.2
11.2
8.2
11.2
Particle
Diameter
ym
0.8
1.6
2.9
0.8
1.6
2.9
0.8
1.6
1.9
0.8
1.6
2.9
(T)
experiment
0.05
0.076
0.27
0.054
0.08
0.306
0.042
0.058
0.336
0.044
0.074
0.362
Jackson §
Calvert
Ci = 10
0.28
1.03
3.23
0.38
1.4
4.4
0.28
1.03
3.23
0.38
1.4
4.4
(NTU/Z) Predicted
Paretsky
0
0
0
0
0
0
0
0
0
0
0
0
Bohm §
Jordan
f '=6.5
2.5
10.0
32.9
3.4
13.7
45.0
2.5
10.0
32.9
3.4
13.7
45.0
Goren
0.
0.
0.
0.
0.
r.
0.
0.
0.
0.
0.
1.
004
07
98
008
15
98
004
07
98
008
15
98
Westinghouse
0
0
0
0
0
0
0
0
0
0
0
0
.078
.14
.25
.106
.19
.34
.78
.14
.25
.106
.19
.34
Schmidt
0.48
0.88
1.54
0.66
1.20
2.10
0.48
0.88
1.54
0.66
1.20
2.10
-------
collection. This does not agree with the experimental data.
Goren's equation underestimates penetration for particles
smaller than 1.6 ym and predicts too low a penetration for
particles larger than 1.6 ym in diameter.
DATA BY PARETSKY, ET AL.
Paretsky et al. (1971) studied the filtration of dilute
aerosols by beds of sand. Test conditions and data were reported
in an earlier section. Their data obtained with a bed of 1,200
to 1,700 ym (-10+14 mesh) angular sand are compared with various
models in Figure 38.
The agreement between Paretsky et al.'s data and theory is
good for gas velocities less than 10 cm/s. For higher gas flow
rates; i.e., in the region where particle collection by impaction
is dominant, the theory underestimates the collection efficiency.
Agreement between Bohm and Jordan's equation and Paretsky
et al's data is poor in the high gas flow region. In the diffu-
sional collection region the agreement is fair.
Predictions by other models do not agree with Paretskyfs
data.
DATA BY GEBHART, ET AL.
Gebhart et al. (1973) published an extensive experimental
study on the collection of aerosol particles by diffusion in
packed beds consisting of uniform glass spheres. They derived
an experimental correlation to predict the diffusional collection
in a granular bed.
Figure 39 shows the predicted diffusional collection in a
granular bed filter by the equation proposed by Paretsky et al.
along with the Gebhart et al. data. The agreement between theory
and data is fair.
Both Goren's equation and B8hm and Jordan's equation predict
much too low a penetration when compared with the Gebhart et al.
data.
116
-------
§
1
»
§
^H
i
PARETSKY ET AL. DATA
UPWARD FLOW
DOWNWARD FLOW
BED DEPTH « 19.2 cm
AEROSOL = 1.1 urn 01A.
PSL
- A. JACKSON AND CALVERT. C
- B. PARETSKY, ET AL.
_ C. BOHM AND JORDAN
0. GOREN
- E. WESTINGHOUSE
F. SCHMIDT, ET AL.
G. GEBHART, ET AL.
1.0
0.05
0.01
5 10
SUPERFICIAL GAS VELOCITY, cm/s
Figure 38. Comparison of Paretsky, et al. data with predictions by
available design equation.
o
o
2
u.
o
g
01
tu
D.
0.1
0.01
0.1
GEBHART
DATA
PARETSKY
ET AL.
PREDICTION^
1.0
dp, urn
Figure 39. Comparison of Gebhart et al. data
and predictions by Paretsky et al.
equation. Collection is in the
diffusion regime.
-------
CONCLUSIONS
Based on the comparisons between theory and data presented
above the following conclusions may be drawn:
1. For superficial gas velocities less than 10 cm/s,
Paretsky's equation can be used to predict granular
bed filter performance.
2. For superficial gas velocities greater than 10 cm/s, the
primary collection mechanism is impaction. The avail-
able models have not been shown to be satisfactory for
the prediction of granular bed collection efficiencies.
118
-------
SECTION 5
EXPERIMENT
APPARATUS
The need for some experimental work became apparent in the
course of the present research. A practical granular bed filter
should be operated at a high gas flow rate where inertial impac-
tion is the principal particle collection mechanism. Available
design equations were not adequate for predicting granular bed
collection efficiencies in the inertial impaction regime.
To obtain further information on the mechanism of particle
collection by impaction and to generate additional clean bed per-
formance data, the experimental apparatus shown in Figure 40 was
constructed. Filtered room air was used for the study and all
flow rates were monitored with rotameters. Monodisperse poly-
styrene latex aerosol was generated using a Collison atomizer.
The aerosol mist from the generator mixed with a stream of dilu-
tion air and passed through a dryer to vaporize the water. Sta-
tic charges were removed by passing the aerosol through a charge
neutralizing section. The charge neutralizing section consisted
of a Krypton-85 charge neutralizer.
Following the neutralizing section the aerosol was further
diluted with filtered room air and then passed into the granular
bed test section. Gas flow through the granular bed was controlled
by using a bypass vent.
The granular bed test section was made of 10.2 cm (4 in.)
I.D. glass pipe and the filter was a bed packed with either iron
shot or sand. The aerosol concentrations before and after the
bed were measured with an optical counter. Pressure drop was
monitored with calibrated gauges.
Five grades of iron shot and one grade of sand were used as
bed materials. The iron shots were SAE S-110, S-170, S-230,
S-280, and S-330. Figure 41 shows the size distributions of
119
-------
VENT
ROTAMETER
OPTICAL
COUNTER
GRANULAR
BED
OPTICAL
COUNTER
PRESSURE
TAP
DILUTION AIR
VENT
ROTAMETER=
FILTER
CHARGE
NEUTRALIZER
DRYER
COLLISON
ATOMIZER
ROTAMETER
FILTER
COMPRESSED
AIR
AIR
BLOWER
Figure 40. Schematic diagram of the experimental apparatus.
120
-------
3,000
I 1,000
o
E
C/3
300
T I T
O s-no
_/\ S-170
S-230
S-330
I
I
I
I
10 20 30 40 50 60 70 80 90
PERCENT BY WEIGHT UNDERSIZE, I
95
98
Figure 41 . Particle size distribution for iron shot.
-------
these shots as measured with sieves. The mass median diameters
are 490 urn, 620 ym, 730 ym, 790 ym, and 860 ym for S-110, S-170,
S-230, S-280, and S-330 shots, respectively. The sand was Agsco
#2 quartz which was obtained from Exxon Research and Engineering
Company. This sand is the same as Exxon used in their Ducon
granular bed filter. The granule size of the sand is -30 +50
mesh. The median diameter is 400 ym.
The experiments were conducted with 0.5 ym, 0.76 ym, and 1.1
ym diameter polystyrene latex monodisperse particles. Figures 42
through 53 show the data. In general, collection efficiency
increases with decreasing granule size, increasing bed depth,
and increasing superficial gas velocity.
DATA ANALYSIS
Table 16 is a list of pressure drops and collection efficien-
cies of 1.1 ym diameter PSL at a superficial gas velocity of 50
cm/s. It reveals that less pressure drop is required for par-
ticle collection with small granules as bed material and shallow
beds instead of deep beds.
For a granular bed with a bed depth of 3.2 cm and operated
at a superficial gas velocity of 50 cm/s, the collection effi-
ciencies for 1.1 ym diameter particles are 22% and 531, respec-
tively, for 620 ym and 490 ym diameter iron shot. The pressure
drop increases from 10 cm W.C. for 620 ym diameter granules to
21 cm W.C. for 490 ym granules. The increase in pressure drop
is 110%. However, by using the finer grade of granules, the
increase in efficiency is 140%.
Table 17 is a list of pressure drops for various beds whose
collection efficiencies are 50% for 1.1 ym diameter particles.
As can be seen the pressure drops for shallow beds are less than
for deep beds.
For a shallow bed to have the same collection efficiency as
a deep bed, it has to run at a high superficial gas velocity.
Therefore, the gas flow capacity of a shallow bed is higher than
that of a deep bed. However, there is a limit on how high a gas
122
-------
1.0
0.9
0.8
S 0.7
| 0.6
S
i i
BED DEPTH = 3.1 cm
S
0.5
0.4
BED MATERIAL: IRON SHOT
GRANULE DIAMETER: 490 \m
AEROSOL: 0.5 urn DIA. POLYSTYRENE LATEX
0.3
10
20
60
30 40 50
SUPERFICIAL GAS VELOCITY, cm/s
Figure 42. Experimental particle penetration of a clean granular bed filter.
70 80 90 100
1.0
0.9
0.8
g 0.7
t—•
§ 0.6
§
t 0.5
0.4
BED DEPTH = 3.2 cm
6.2 cm —
BED MATERIAL: IRON SHOT
GRANULE DIAMETER: 620 pm
AEROSOL: 0.5 pm DIA. POLYSTYRENE LATEX
0.3
i
1
1
i
10 20 30 40 50 60 70 80 90 100
SUPERFICIAL GAS VELOCITY, cm/s
Figure 43. Experimental particle penetration of a clean granular bed filter.
123
-------
1.0
0.9
0.8
2 °'7
t—
£ 0.6
o
i 0.5
h—
UJ
z
UJ
ex
0.4
0.3
BED DEPTft =3.2 cm
6.2 cm
9.2 cm
12.2 cm
BED MATERIAL: IRON SHOT
GRANULE DIAMETER: 790 ,.m
AEROSOL: 0.5 ..m DIA. POLYSTYRENE LATEX
10
20 30 40 50
SUPERFICIAL GAS VELOCITY, cm/s
60
70 80 90 100
Figure 44. Experimental particle penetration of a clean granular bed filter.
1.0
0.9
0.8
0.7
0.6
0.5
0.4
0.3
BED DEPTH 6.2 cm
BED MATERIAL: IRON SHOT
GRANULE DIAMETER: 620 i.m
AEROSOL: 0.76 ;.m DIA. POLYSTYRENE LATEX
10
60
70 80 90 100
Figure 45. Experimental particle penetration of a clean granular bed filter.
20 30 40 50
SUPERFICIAL GAS VELOCITY, cm/s
124
-------
1.0
0.9
0.8
S 0.7
M
3 0.6
2 0.5
tu
0.4
0.3
BED DEPTH = 9.2 cm
12.2 cm
BED MATERIAL: IRON SHOT
GRANULE DIAMETER: 790 urn
AEROSOL: 0.76 \an DIA. POLYSTYRENE LATEX
10
20 30 40 50
SUPERFICIAL GAS VELOCITY, cm/s
60 70 80 90 100
Figure 46. Experimental particle penetration of a clean granular bed filter.
1.0
0.9
0.8
0.7
0.6
0.5
0.4
0.3
0.2
9.2 cm
12.2 cm
BED MATERIAL: IRON SHOT
GRANULE DIAMETER: 620 pm
AEROSOL: 1.1 pm DIA.
POLYSTYRENE LATEX
20 30 40 50 60 70 80 90 100
SUPERFICIAL GAS VELOCITY, cm/s
Figure 47. Experimental particle penetration of a clean
granular bed filter.
125
-------
1.0
0.9
0.8
0.7
0.6
0.4
0.3
0.2
BED DEPTH =
3.2 cm
12.2 cm
BED MATERIAL: IRON SHOT
GRANULE DIAMETER: 730 urn
AEROSOL: 1.1 ym DIA. POLYSTYRENE
LATEX
20 30 40 50 60 70 80 90 100
SUPERFICIAL GAS VELOCITY, cm/s
Figure 48. Experimental particle penetration of a clean granular
bed filter.
1.0
0.9
0.8
0.7
0.6
P 0.5
2 0.4
0.3
0.2
BED DEPTH =
3.2 cm
6.2 cm
9.2 cm
12.2 era
BED MATERIAL: IRON SHOT
GRANULE DIAMETER: 860 \m
- AEROSOL: 1.1 urn DIA. POLYSTYRENE LATEX
20
30 40 50 60 70 80 90 100
SUPERFICIAL GAS VELOCITY, cm/s
Figure 49. Experimental particle penetration of a clean granular
bed filter.
126
-------
1.0
IS)
1.0
0.9
0.8
0.7
§ 0.6
i—*
(J
U_
„ 0.5
<£
ce.
0.4
0.3
0.2
BED DEPTH =
3.2 cm
6.2 cm
9.2 cm
12.2 cm
BED MATERIAL: IRON SHOT
GRANULE DIAMETER: 790 urn
AEROSOL: 1.1 urn DIA. POLYSTYRENE LATEX
20 30 ' 40 CO 60 70 80 90 100
SUPERFICIAL GAS VELOCITY, cm/s
Figure 50. Experimental particle penetration of a clean granular
bed filter.
§
1
0.5
0.4
0.3
0.2
0.1
0.05
9.2 cm
BED MATERIAL: IRON
GRANULE DIAMETER: 490
AEROSOL: 1.1 um DIA
SHOT
\im
PSL
I
I
10
I I
20 30 40 50
SUPERFICIAL GAS VELOCITY, cm/s
100
Figure 51. Experimental particle penetration of a clean granular
bed filter.
-------
1.0
is}
oo
0.5
0.2
0.1
0.05
0.04
0.03
1IIIITT
0.5 urn DIA. PSL
— BED MATERIAL: AGSCO SAND
— GRANULAR DIAMETER: -30+50
MESH
- BED DEPTH: 1.7 cm
I
I I
10
20 30 40 50
SUPERFICIAL GAS VELOCITY, cm/s
100
Figure 52. Experimental particle penetration of a clean granular
bed filter.
1.0
0.5
0.1
o
r 0.05
0.01
0.05
0.03
I I I I I I I <-
0.5 pro DIA. PSL
1.1 pm DIA. PSL
BED MATERIAL: AGSCO SAND
GRANULE DIAMETER: -30+50 MESH
BED DEPTH: 2.6 cm
I III JIM
I I I I I I I I
5 10
SUPERFICIAL GAS VELOCITY, cm/s
50 100
Figure 53. Experimental penetration of a clean granular bed
filter.
-------
TABLE 16. PRESSURE DROP AND COLLECTION EFFICIENCY FOR 1.09 urn
DIAMETER PARTICLES AT A SUPERFICIAL GAS VELOCITY OF 50 cm/s
Granule
Diameter
(ym)
490
620
730
790
860
Bed
%
Coll
53
22
14
12
10
Depth = 3.2 cm
AP
(cm W.C.)
21
10
7
65
6
Bed
%
Coll
83
38
27
30
19
Depth = 6.2 cm
AP
(cm W.C.)
38
20
14
12
12
Bed
%
Coll
93
55
35
44
27
Depth = 9.2 cm
AP
(cm W.C.)
_-
28
21
18
28
Bed
%
Coll
_-
68
53
55
38
Depth » 12.2 cm
AP
. (cm W.C.)
—
37
27
23
23
to
-------
TABLE 17. PRESSURE DROP FOR 50% COLLECTION
OF 1.1 urn DIAMETER PARTICLES
Granule
Di ameter
(ym)
490
620
730
790
860
Pressure Drop (cm W.C.)
Bed Depth
3.2 cm
19
17
16
15
15
Bed Depth
6.2 cm
--
24
23
20.5
23
Bed Depth
9.2 cm
--
26
27
22
28
Bed Depth
12.2 cm
—
22
25
21
30
130
-------
velocity the bed can be safely operated without the danger of
causing particle reentrainment.
PRESSURE DROP DATA
The pressure drop data obtained with the iron shot bed mate-
rial were analyzed. Figures 54 through 58 show the experimental
and predicted pressure drops. The prediction was based on
Ergun's equation. As can be seen, the predicted pressure drop
is lower than that measured.
The pressure drop prediction based on Ergun's equation is
very sensitive to the bed porosity. The difference in the pre-
dicted and measured pressure drops might be caused by an error
in bed porosity determination. The bed porosity was calculated
from the measured weight of a bed of known volume. It is very
difficult to accurately determine bed porosity by this method.
By fitting Ergun's equation to the pressure drop data, the
iroid fraction of the bed was obtained. Table 18 shows the result
along with the measured void fractions. The measured void frac-
tion is very close to calculated void fraction.
TABLE 18. MEASURED AND CALCULATED VOID FRACTION OF
THE GRANULAR BED
Shot No.
S-110
S-170
S-230
S-280
S-330
Average Diameter
Vim
490
620
730
790
860
Measured
Void Fraction
0.39
0.39
0.41
0.41
0.40
Void Fraction From
Pressure Drop Data
0.37
0.39
0.40
0.40
0.39
131
-------
100
to
100
50
§ 40
a.
i
° 30
ae
20
10
I I \
PREDICTED FOR e = 0.39
I I I II
BED DEPTH =
9.2 cm
BED MATERIAL: IRON SHOT
GRANULE DIAMETER: 490
I
I I I I
20 • 30 40 50 100
SUPERFICIAL GAS VELOCITY, cm/s
Figure 54. Experimental and predicted pressure drops across a
clean granular bed filter.
o
on
o
50
40
30
20
10
T
I
I I I I I
PREDICTED FOR e = 0.39
BED MATERIAL: IRON SHOT
GRANULE DIAMETER: 620 urn
10
BED DEPTH
12.2 cm~*
9.2 or-
6.2 cm-
3.2 cm
I
I I I I I
20 30 40 50
SUPERFICIAL GAS VELOCITY, cm/s
100
Figure 55. Experimental and predicted pressure drops across a
clean granular bed filter.
-------
CM
50
40 I—
30
20
,o
1 I
PREDICTED FOR e = 0.41
BED MATERIAL: IRON SHOT
GRANULE DIAMETER: 730 urn
10
BED DEPTH
12.2 cm
9.2 cm
5.2 cm —'
3.2 cm ^*
I
20 30 40 50
SUPERFICIAL GAS VELOCITY, cm/s
100
Figure 56. Experimental and predicted pressure drops of a clean
granular bed filter.
O
a:
a
in
UJ
ce.
50
40
30
20
10
I I
PREDICTED FOR
BED MATERIAL: IRON SHOT
GRANULE DIAMETER: 790 MI
I I
6 = 0.41
10
BED DEPTH
12.2 cm-
3.2 cm
I
I
I
I I I I
20 30 40 50
SUPERFICIAL GAS VELOCITY, cm/s
100
Figure 57. Experimental and predicted pressure drops across a
clean granular bed filter.
-------
40
T I I I I I
PREDICTED FOR e = 0.40
30
BED MATERIAL: IRON SHOT
GRANULE DIAMETER: 860 ym
20
o
CL.
O
UJ
10
I I I I I I
10 20 30 40 50 100
SUPERFICIAL GAS VELOCITY, cm/s .
Figure 58. Experimental and predicted pressure drops across a
clean granular bed filter.
134
-------
SECTION 6
DESIGN MODEL
MATHEMATICAL MODELING
The granular bed can be envisioned as a great number of
impaction stages connected in series (Figure 59). Particle
collection is by impaction as in a cascade impactor. The jet
openings are the pores in each layer of granules. It is assumed
that the jet diameters in the granular bed are of uniform size
with a diameter equal to the hydraulic diameter of the void
space. The gas velocity in the jet is the average interstitial
gas velocity.
If 'n* is the collection efficiency of one impaction stage,
the particle penetration for the granular bed will be
Ptd - (l-n)N (88)
where Ft, = penetration for particles with
diameter d_, fraction
n = single stage collection efficiency, fraction
N = number of impaction stages, number
As in some cascade impactors, each layer of granules served
both as the jet plate and as the collection plate. Therefore,
each layer of granules is an impaction stage and "N" is equal to
the number of granular layers in a bed. For a randomly packed
bed
N = 1 1 C89)
2dc
where Z- bed depth, cm
d = granule diameter, cm
135
-------
jet formed by
granule cluster
granule
target
Figure 59. Diagram of granular bed showing
impaction concept.
136
-------
and
Ptd - U-n) ' dc" (90)
The impaction collection efficiency, "n" is a function of
"Kp", the inertial impaction parameter. The impaction parameter
is defined as
v C' Pp dp uj
P Q 11 A L (91}
c y y^ a.
u i
where c' = Cunningham slip factor, dimensionless
Pp = particle density, g/cm3
dp = particle diameter, cm
u. = jet velocity, cm/s
UG » gas viscosity, g/cm-s
d. = jet diameter, cm
Since U;j = uQi = ^ (92)
and d. = 4rH = | ^- dc (93)
where uGi = average interstitial gas velocity,
cm/s
e = bed porosity, fraction
rH = hydraulic radius, cm
d = granule diameter, cm
we have Kp - | if C9PgdP "G (94)
£ bC
The relation between "n" and "K " can be evaluated once the
flow field is defined. Flow fields reported in the literature for
inertial impaction, e.g., Ranz and Wong (1952) and Marple (1970),
are adequate for K > 0.15. For Kp < 0.15, there is no suitable
flow field reported in the literature. Therefore, the relationship
between "n" and "K " could not be calculated analytically.
137
-------
The relationship was back-calculated from equation (90) and
experimental data. Figure 60 shows the results. The curve can
be approximated by the following equation:
n = 10.0 K 3'23 exp [0.27 (In Kp)2] (95)
for 3 x 10"3 < K < 0.15
P -
There is scatter in the lower end of the curve. For K < 10" ,
"n" is very sensitive to experimental data. A few percent scatter
in the data will cause "n" to fluctuate greatly. Figure 61 com-
pares the experimentally determined "n" versus "K " curve with
those reported by Ranz and Wong (1952), Stern, et al. (1962),
and Mercer and Stafford (1969). All reported curves are for
K > 0.15. As can be seen, the curve calculated in the present
study matches other researchers' results. The curve determined
in this study is a continuation of other researchers' curves.
Paretsky et al. (1971) and Knettig and Beeckmans (1974)
studied the collection of monodispersed aerosol particles in
granular bed filters. Their data were transformed into "K "
versus "n" plots as shown in Figure 62.
Knettig and Beeckmans used 425 ym glass beads as granular
material. Bed porosity was 0.38. Aerosol particles were 0.8,
1.6, and 2.9 ym in diameter. As can be seen from Figure 62,
their data are close to the results of present study.
Paretsky et al. (1971) studied the filtration of 1.1 ym
diameter polystyrene latex aerosols by beds of sand. They studied
a bed of -10+14 mesh (1,200 to 1,700 ym) angular sand and a bed
of -20+30 mesh (500 to 850 ym) sand at superficial gas velocities
between 0.3 and 80 cm/s. Bed porosities were 0.41 and 0.43,
respectively. Single stage collection efficiencies were calcu-
lated from their data. In the calculation, the granule diameters
were assumed to be the arithmetical mean of the smallest and the
largest granule size in the bed. The results are plotted in
Figure 62. For a given inertial parameter, Paretksy et al.
138
-------
0.1
§
hH
H
U
u,
10
-2
10
-3
10
-4
1.09 vim
0.76 ym
0.5 urn
3*2 3 2
n = 10 Kp exp [0.27 (In K ) ]
2x10
10
0.1
K , DIMENSIONLESS
Figure 60.
K versus n for round jet model for
particle collection in a GBF.
139
-------
§
CJ
i.o
0.5
0.1
o.os
0.01
0.005
0.001
10
STERN
ET AL,
(1962)
RANZ §
WONG
(1952)
MERCER §
STAFFORD
(1969)
—- PRESENT STUDY
O.OS
0.1
0,5
K , DIMENSIONLESS
0.05
10
LJ
10"
10
I I I I I '
o
KNr.TTlC. f,
Bl:.r.i:KMANS
(1974)
PARHTSKY I:.T AL./
20-30 MESR,
. PARMTSKY F.T AL.
11971) X
10-14 MESH X-
PRESENT
STUDY
i i i i
i i i i i i
0.002
0.01
K DIMENSIONLESS
P.
0.1
0.2
Figure 61, 'K ' versus 'n' for round jet.
Figure 62. Efficiency vs. inertial impaction
parameter for comparison.
-------
data give a higher collection efficiency than reported in this
study. Their data would be close to that of present study if
the mean diameter were equal to the smallest granule diameter.
COMPARING MODEL PREDICTIONS WITH PERFORMANCE DATA
The design equation is for the prediction of particle col-
lection by a clean bed. If no filter cake is formed and
the collected particles are uniformly distributed in the bed, then
the equation is still applicable. The design equation has been
used to predict the performance of Rexnord gravel bed filters
and the Combustion Power Company "dry scrubber."
Figure 63 shows the comparison between the data reported by
McCain (1976) for a Rexnord gravel bed filter and the prediction by
the present model. The prediction by the present model is close
to the data.
Figure 64 shows the comparison of Hood's data with predic-
tions. The predicted penetration is higher than that measured.
Figures 65 through 67 show the comparison of A.P.T.'s data
for CPC GBF with predictions. The present model predictions
agree with data for some runs but not for all runs. The present
model is very sensitive to bed porosity. In the calculations,
a bed porosity of 0.35 was used. If a bed porosity of 0.3 is
used, the model would give a better fit with the data.
141
-------
1.0
0.5
§ 0.1
I
2
£ 0.05
0.01
I I
A. JACKSON & CALVERT
B. PARETSKY, ET AL.
C. BbHM & JORDAN,
f = 10
D. B'OHM AND JORDAN,
f = 6.5
E. GOREN.e = 0.25
F. WESTINGHOUSE
6. SCHMIDT, ET AL.e = 0.25
D
I I I I I I I I !
0.1
0.5 1
AERODYNAMIC PARTICLE DIAMETER, umA
10
Figure 63. Comparison of McCain's gravel bed particle collection data
with design equation predictions.
1.0
0.5
2
i
0.05
0.01
i i i r
HOOD'S
DATA —
A. JACKSON & CALVERT
B. PARETSKY, ET AL.
C. BOHM & JORDAN,
f'= 10
D. BOHM & JORDAN, f=6.5
I I I I I I I I I
E. GOREN.e = 0.25
F. WESTINGHOUSE
G. SCHMIDT, e = 0.25
I I I I I I I I
0.1
0.5 1
PARTICLE DIAMETER, urn
10
Figure 64. Comparison of Hood's data with predictions by available
design equations.
-------
data give a higher collection efficiency than reported in this
study. Their data would be close to that of present study if
the mean diameter were equal to the smallest granule diameter.
COMPARING MODEL PREDICTIONS WITH PERFORMANCE DATA
The design equation is for the prediction of particle col-
lection by a clean bed. If no filter cake is formed and
the collected particles are uniformly distributed in the bed, then
the equation is still applicable. The design equation has been
used to predict the performance of Rexnord gravel bed filters
and the Combustion Power Company "dry scrubber."
Figure 63 shows the comparison between the data reported by
McCain (1976) for a Rexnord gravel bed filter and the prediction by
the present model. The prediction by the present model is close
to the data.
Figure 64 shows the comparison of Hood's data with predic-
tions. The predicted penetration is higher than that measured.
Figures 65 through 67 show the comparison of A.P.T.'s data
for CPC GBF with predictions. The present model predictions
agree with data for some runs but not for all runs. The present
model is very sensitive to bed porosity. In the calculations,
a bed porosity of 0.35 was used. If a bed porosity of 0.3 is
used, the model would give a better fit with the data.
141
-------
IX)
1.0
0.5
g
i
0.1
0.05
0.01
T I Tl
— A. JACKSON & CALVERT
- B. PARETSKY, ET AL.
"" C. BbHM & JORDAN,
f = 10
D. B'OHM AND JORDAN,
f = 6.5
- E. SOREN.e = 0.25
F. WESTINGHOUSE
G. SCHMIDT, ET AL.e = 0.25
D
I I I I I I I I
I I i I I I J
0.1
0.5 1
AERODYNAMIC PARTICLE DIAMETER, umA
10
Figure 63. Comparison of McCain's gravel bed particle collection data
with design equation predictions.
1.0
0.5
0.1
0.05
0.01
A. JACKSON & CALVERT
B. PARETSKY, ET AL.
C. BOHM & JORDAN,
f = 10
D. BOHM & JORDAN, f=6.5
I I I I I I
E. GOREN.e = 0.25
F. WESTINGHOUSE
G. SCHMIDT, e = 0.25
I I
I I
I I I I I
o.i
0.5 1
PARTICLE DIAMETER,
10
Figure 64. Comparison of Hood's data with predictions by available
design equations.
-------
w
1.0
0.5
0.1
UJ
z
LU
0.05
0.01
UQ = 81 cm/s
dc = 0.2 cm
Z = 20.2 cm
GRANULE CIRCULATION
RATE = 0.8 kg/kg AIR
A.P.T. DATA
SCHMIDT. ET AL. .
e = 0.35 \
J I I I I I I I
I I I I I I I
0.1
0.5 1
AERODYNAMIC PARTICLE DIAMETER,
10
Figure 65. Experimental and predicted performance of CPC GBF
(A.P.T. data).
1.0
0.5
, 0.1
fe 0.05
0.01
1 I
PRESENT MODEL,
E = 0.35
WESTINGHOUSE
0.1
SCHMIDT, ET AL
t. = 0.35
UG = 41 cm/s
dc = 0.2 cm
A = 20.2 cm
I I I I I I
GOREN, E = 0.35 _|
GRANULE CIRCULATIOirl
RATE = 1.6 kg/kg -\
/ AIR
A.P.T. DATA
GRANULE RECIRCULATION \
RATE = 0.8 kg/kg AIR
I I I I I I I I I
I I
I I
0.5 1
AERODYNAMIC PARTICLE DIAMETER,
10
Figure 66. Experimental and predicted performance of CPC GPF
(A.P.T. data).
-------
1.0
0.5
I I I I I I I I I
PRESENT MODEL
e = 0.35
WESTINGHOUSE-
o
i—
o
2
oi
LU
LU
Q-
0.05
SCHMIDT, ET AL.
e = 0.35
UG =
I" dc =
I =
41 cm/s
0.2 cm
40.6 cm
I I I I I I I
GOREN, e = 0.35 _
0.01
I I I I I I I I I
\l I I I I I I I
0.1 0.5 1 5
AERODYNAMIC PARTICLE DIAMETER, ymA
Figure 67. Experimental and predicted performance of CPC GBF
(A.P.T. data).
10
144
-------
SECTION 7
PRESENT TECHNOLOGY EVALUATION
GRANULAR BED FILTER SYSTEMS
GBFs have been used commercially for over 30 years, with
several designs commercially available. Granular bed filters
may be classified either according to bed structure or according
to bed cleaning method. The first classification method is used
in this report.
With respect to the bed structure, granular bed filters
may be classified as continuously moving, intermittently moving,
and fixed bed filters.
Continuously Moving Bed Filters
The continuously moving bed filter is usually arranged in a
cross-flow configuration. The bed is a vertical layer of granular
material held in place by louvered walls. The gas passes hori-
zontally through the granular layer while the granules and col-
lected dust continuously move downward and are removed from the
bottom. The dust and granules are separated by vibration. The
cleaned granules are then returned to the overhead hopper by a
granule circulation system.
Several commercial designs fall into this category. They
include the Dorfan Impingo filter, the Consolidation Coal Company
filter, and the "dry scrubber" of the Combustion Power Company.
The "dry scrubber" is the only one that is presently marketed.
Dorfan Impingo Filter -
The Dorfan Impingo filter was invented by Morton Dorfan and
was offered commercially by Mechanical Industries, Inc. in the
early 1950's. The device is a vertical panel filter in which the
granular materials continuously fall through the panel. The
granule flow rate is controlled by the setting of a rotary valve
145
-------
located near the bottom of the panel. The dust-laden gas is
filtered by blowing the gas through the panel horizontally.
Filtered dust is carried downward with the granules (Figure 68).
The opposite walls of the panel are not parallel but are
slightly offset from the vertical so that the panel is tapered
with its narrowest point at the top and widest at the bottom.
The granules thus travel downward in a panel of constantly in-
creasing cross-sectional area. The rationale is that the tapered
construction acts to prevent hangups caused by size increases of
the granules as the dust loading builds upon their surfaces.
Four units, each of 8 m3/s (17,000 CFM) gas capacity and
consisting of two cells, were installed in a plant to collect
asbestos rock dust from a stream of flue gas coming from a direct
fired dryer in which the rock was dried prior to milling. The
granular bed was a 30 cm (1 ft) thick panel with 2.74 m x 4.27 m
(9 ft x 14 ft) filtering area. The granules were raw asbestos
rock ranging from 1.3 cm to 3.8 cm (0.5 to 1.5 in.) in diameter.
The dust was 100% finer than 100 mesh and 601 finer than
10 microns. The concentration entering the collector was approxi-
mately 14.8 g/m3 (6 g/ft3) and that leaving about 0.49 g/m3 (0.2
g/ft3).
The Dorfan Impingo filter installations are no longer in
use and the equipment is not presently marketed.
Consolidation Coal Company Filter -
A granular bed filter was studied first on a pilot scale
and then with a full-size installation by the Consolidation Coal
Company in the period 1950-1952. This equipment was used for
collecting dust from hot gases leaving coal drying operations.
Lump coal [0.95 cm to 3.2 cm (3/8 in. to 1.25 in.)in size] was
used as the bed granules. The design concept for the full size
installation is shown in Figure 69.
The granules were fed from an overhead bin to two separate
panels. The panel at each side consists of six separate cells
facing the inlet gas which flowed through the beds from a central
dust and exited via ducts at either side of the installation.
Typical operating characteristics are given in Table 19.
146
-------
Filter
Particles
Rotary Valve
Figure 68. Dorfan Impinge Filter.
147
-------
Granule Feed
Granular Discharge
Figure 69. Consolidation Coal Company filter
148
-------
TABLE 19. TYPICAL OPERATING CHARACTERISTICS OF
CONSOLIDATION COAL COMPANY GRANULAR
BED FILTERS
Type of service - recovery of dust from coal drying operation,
Size - Individual cell was 127 cm wide x 122 cm high x 51 cm
thick (4'2" wide x 4' high x 20" thick). In the unit
the filtering surfaces were divided into two banks of
six cells each in parallel.
Capacity - 0.5 - 1.5 m3/s per square meter of filter area
(100 - 300 CFM/ft2)
Total filter area - 18.6 m2 (200 ft2)
Inlet dust loading - 7.4 - 17.2 g/m3 (3 - 7 gr/SCF)
Size of dust - 86% < 30 ym and 22% < 10 pm
Granular material - Coal, 0.95 cm to 3.2 cm (3/8" to 1-1/4")
Granule flow rate - 22 metric tons/hr
Operating temperature - 130°F
Pressure drop - 5.8 cm W.C.
Collection efficiency - 89 - 99%
149
-------
Combustion Power Company "Dry Scrubber" -
The "Dry Scrubber" is currently offered commercially by Com-
bustion Power Company. It is similar to the Dorfan Impingo Filter.
It consists of a vertical bed of granules held in place by louvers.
It is operated in cross-flow configuration. Gas flows horizontally
through the bed and the bed continuously moves downward. Clean
granules are introduced at the top and a mixture of dust and
granules is removed from the bottom. Dusts and granules are
separated in a shaking device.
Two models are produced by Combustion Power Company. Figure
70 shows the regular "Dry Scrubber." It is used when the parti-
culate loading is low. For very high inlet particle loadings, the
"Integral Cyclone Model" is recommended. It has a low energy
cyclone wrapped around the outer shell of the standard "Dry Scrub-
ber" (Figure 71].
Intermittently Moving Bed Filters
In the late 1950s, Squires modified the continuously moving
bed design of the Dorfan filter to obtain a fixed bed device with
intermittent movement of granular solids. The design is called
the LS (Loose Surface) filter. It uses a finer grade granule than
the Dorfan filter and the bed is stationary during filtration.
The accumulated filter cake is removed by moving just the surface
layer of granules.
Figure 72 shows one possible arrangement. The granular bed
is a narrow vertical bed of granules (-50+60 mesh) held between a
panel of louvers and a fine mesh screen. In some applications,
a relatively coarse grade granule is used on the gas exit side
to prevent blow-through of the smaller collecting granules. During
operation the dusty gas flows horizontally through the panel bed.
The particles collected by the bed build up a cake on the exposed
bed surfaces and to some extent penetrate to the interstices.
When the resistance of the cake has reached an undesirable level,
the clean gas outlet valve is closed and a short pulse of compressed
air is blasted in reverse flow through the granular bed. In con-
tinuous use, the valves operate on a timed cycle.
The blow-back pulse is sufficient to physically lift the
150
-------
GAS
FLOW
CLEAN
MEDIA
CONVEYOR
TO DUST STORAGE
Figure 70. Combustion Power Company "Dry Scrubber."
151
-------
CLEAN
MEDIA
CONVEYOR
TO DUST STORAGE
Figure 71. The intergral cyclone model of the "Dry Scrubber."
152
-------
DIRTY GAS
)IRTYI GAS
CROSS-SECTION
OF COLUMN
CLEAN GAS,,
OUT
BLOWBACK
VALVE
SAND BED
DIRTY
GAS
CLEAN
GAS
CROSS-SECTION
OF FILTER BED
SAND REMOVAL CHUTE
Figure 72. Possible design for Squires Panel Bed Filter.
153
-------
sand beds as a mass, with minimum inter-particle movement, so that
a surface layer of granules between each pair of louvers is physi-
cally ejected from the panel and falls to the bottom of the filter
vessel along with the collected filter cake. The expelled granule
is immediately replaced by the downward movement of a fresh granule
from the overhead hoppers.
The development of the LS filter has continued for the past
ten years at the City College of the City University of New York
with financial support from EPA and EPRI. This development work
resulted in minor modifications of louver configurations (wish-
bone type louvers) and of the puff-back cleaning technique.
Fixed Bed Granular Filters
As opposed to the continuous and intermittent moving beds,
fixed bed granular filters require no granule circulation. Col-
lected particles in the bed are removed either mechanically or
penumatically. There are three fixed bed devices. The "Lurgi-
MB Filter" and the "Rex-Gravel Bed Filter" clean the bed mechani-
cally. The "Ducon Granular Bed Filter" uses a reversed gas flow
to clean the bed.
The Lurgi MB Filter -
Max and Wolfgang Berz designed a granular bed that requires
no granule circulation. The cleaning is carried out by flowing
a reverse flow of gas through the bed while subjecting it to
mechanical vibration of sufficient magnitude to cause the inter-
granule movement necessary for removal of entrapped particles.
The granular bed is a layer of loosely packed material
such as gravel held on a horizontal sieve plate. Gravel sizes
range from 1 to 6 mm in diameter. Figure 73 is a sketch of the
system.
In operation, the dusty gas flows upward through the gravel
bed and vents through the clean gas duct on top. The dust will
be retained and gradually will buildup in the bed. This increases
the pressure drop. As soon as the pressure drop of the filter
exceeds a pre-determined level, the gas flow is stopped and the
bed is cleaned. For some applications it is possible to stack two
154
-------
-i
Inlet for dirty gas
Gravel layer
Shaking motor
Shut-off flap
Scavenging air flap
Clean gas outlet
Hopper
Screw-type discharger
Figure 73. Lurgi MB filter.
155
-------
or more filter beds, one over the other, in a given device, with
the bottom-most layer consisting of relatively coarse material
to act as a pre-filter.
This granular bed filter was marketed commercially by Lurgi
Appareteban Gesellschaft M.B.H. of Frankfurt, Germany. It was
mostly used in cement plants. It was redesigned in 1968 and was
withdrawn from the market at about the end of 1969. Its short-
coming lies in the strain imposed on the necessary flexing membranes
associated with the vibrating technique. It is conceivable that
at low temperatures where rubber membranes and spring-supported
bed mounts are feasible, such a filter could operate with a rea-
sonable life. However, at low temperatures it could not compete
economically with baghouses. At high temperatures, metal bellows
would be needed for the flexing membrane, and their life expectancy
in the hot and dusty environment is too short for practical
application.
Rex Gravel Bed Filter -
Berz designed an alternate arrangement of the "Lurgi-MB
Filter" and marketed it through Gesellschaft fur Enstaubungsau-
lagen (GfE) of Munich, Germany. In the U.S. it is built and
marketed by Rexnord in accordance with an exclusive license agree-
ment with GfE. It is built in modules for the treatment of large
gas volumes, with several modules arranged in parallel through a
common raw gas duct and a common clean gas duct.
Each gravel bed filter module consists of a filter top
section containing two horizontal beds connected in parallel,
and a cyclone pre-cleaner located underneath. The operation of
the system is illustrated in Figure 74. The raw gas enters the
filter through an inlet chamber where immediate separation (set-
tling) of very coarse materials takes place. From there, the gas
enters the cyclone separator where more coarse dust is separated
and removed through the discharge airlock at the outlet.
The gas then rises from the cyclone through the vortex tube
and enters the filter chambers. It passes from the top of the
horizontal filter beds to the bottom, so that the remaining fine
156
-------
OPERATING PHASE
BACKFLUSH PHASE
12
14
BACKFLUSH
AIR
1. INLET CHAMBER
2. PRIMARY COLLECTOR (CYCLONE)
3. DOUBLE TIPPING GATE (DUST DISCHARGE)
4. VORTEX TUBE
5. FILTER CHAMBER
6. GRAVEL BED
7. SCREEN SUPPORT FOR BED
8. CLEAN GAS COLLECTION CHAMBER
9. EXHAUST PORT
10. BACKWASH CONTROL VALVE
11. BACKWASH AIR INLET
12. VALVE CYLINDER
13. STIRRING RAKE
14. STIRRING RAKE MOTOR/REDUCERS
Figure "74. Rexnord gravel bed filter.
157
-------
dust is deposited on the quartz grains and in the interstices of
the bed. The cleaned gas flows through the clean gas collection
chamber and passes through the 3-way valve into the clean gas
duct.
Cleaning of the filter unit can be initiated by means of a
pre-set sequencer or by automatically monitoring the pressure
differential across the filter bed. During the cleaning cycle,
the unit is isolated from the gas stream by the 3-way valve. Then
backwash air is admitted to the filter chamber in a reverse flow
direction. It is either forced in by using a backwash air blower
or sucked in by negative pressure. The backwash air loosens the
filter bed.
During the cleaning process the rake-shaped double arm stir-
ring device is rotated by the geared motor. This helps the dust
to be removed from the gravel and entrained by the backwash cleaning
operation. The large agglomerated dust particles are carried
by the backwash air via the vortex tube into the precipitator,
where the velocity is reduced and the gas stream deflected so
that a large percentage of the dust is settled out. The backwash
air, containing the remaining dust, mixes with the dust-laden
air in the raw gas duct and is then subjected to cleaning in the
remaining units of the filter.
Ducon Granular Bed Filter -
The Ducon granular bed filter (Figure 75) consists of multiple
beds of sand stacked vertically within two perforated concentric
metal tubes. The beds of sand grains rest on slotted inner screen
supports having opening dimensions slightly smaller than the dia-
meter of sand grains. Above each bed there is an annular space
with an outer screen similar to the inner screen. Thus, each
sand bed contained within the two concentric cylinders is supported
and caged by slotted metal retaining screens. Figure 76 shows a
typical filter element. The elements are supported from a clean
gas plenum within a housing. The basic arrangement is exemplified
in the cutaway sketch of a typical 4-element unit shown in Figure 75.
158
-------
FILTER
ELEMENT
INLET
CLEAN
GAS
OUTLET
COLLECTED
DUST
OUTLET
Figure 75. Ducon granular bed filter.
159
-------
OUTER
SCREEN
GRANULAR
BED
INNER SCREEN
Figure 76. Filter element.
160
-------
In operation, particle laden gas entering the dusty side of
the filter housing passes through the element's outer retaining
screens, flows downward through the inner cylinder into the clean
gas plenum to the vessel outlet nozzle. The particles are deposited
on the surface and within the interstices of the sand beds. To
clean the element, a small volume of compressed air is introduced
in a reverse pulse which induces blowback gas flow from the clean
gas plenum, sufficient to momentarily fluidize all the sand beds
within the element simultaneously. This blow-back gas flexes the
bed in fluidized expansion, expelling interstitially deposited
dusts, as well as any dusts on the bed surfaces, through the outer
screen. The dust then falls into the collector hopper for eventual
removal. Figures 77 and 78 illustrate the collection and cleaning
cycles.
INDUSTRIAL USERS AND PERFORMANCES
Granular bed filters have been used in recent years on selected
sources. While the use of granular beds has not been directed
at the control of fine particulates, granular beds have been used
successfully on cement and lime kilns, asphalt dryers, and clinker
coolers. Of the three commercially available granular bed filter
systems, only the Rexnord granular bed filter and the Combustion
Power Company's moving bed granular bed filter have industrial
installations. Ducon granular bed filters have been used in some
pilot studies.
We surveyed industrial granular bed filter users to obtain
performance data and to identify operating problems. The following
is a summary of the results of this survey.
Rexnord Granular Bed Filter
There are over thirty industrial users of the Rexnord granu-
lar bed filters. Most of the installations are in Portland cement
plants to control particulate emissions from clinker coolers.
The average size of the dust particles from clinker coolers is
relatively large, and the granular bed adequately meets the emis-
sion standards. However plugging of the granule retaining screens
has occasionally caused problems.
161
-------
DUST
LADEN
GAS
— ,
CLEAN
GAS
Figure 77. Collection cycle.
COLLECTED
DUST 4
•• v
A
A
.
r-'i
PURGE
GAS
LUIDIZED
GRANULES
Figure 78. Cleaning Cycle
162
-------
There is one unit which is installed in a steel sinter plant
to control the emissions from a windbox exhaust. The installation
consists of 24 modules with a total capacity of 4,400 Nm3/min
(240,000 ACFM @ 300°F). The bed is packed with 0.1 cm diameter
gravel to a depth of 9 cm. The unit was started up in February
1976 and encountered several operational difficulties. Design
modifications were required. Among the operational difficulties
encountered were:
1. The gravel medium developed growth problems; i.e.,
collected dust adhered to granules and could not be removed.
2. Downcomer valves for dirty gas had to be modified.
3. Backflush fans were replaced by larger fans.
4. Heavier gauge bed support screens had to be substituted
for the originals.
5. Backflush air preheater burners have malfunctioned
repeatedly.
Some users supplied cost data. Table 20 summarizes the cost
data available.
McCain (1976) conducted a performance test on a Rexnord
granular bed installed in a Portland cement plant. Samples were
taken simultaneously at the filter inlet and outlet with cascade
impactors. Particle size distributions and grade efficiencies
were calculated from the impactor data.
The particles were found to have a mass median diameter of
about 200 ymA. The overall collection efficiency was found to
be from 99.3 to 99.7%. The system pressure drop ranged from 9.6
to 14 cm W.C. (3.8 to 5.5 in. W.C.). The system energy usage
during the tests was approximately 1,780 joules/Nm3 (47.7 in W.C.).
Figure 79 shows the grade efficiency curves for three sampling
runs reported by McCain for three cascade impactor runs. Table 21
lists the operational conditions for the granular bed filter
during the test period.
Combustion Power Company's Moving Bed Granular Bed Filter
Four units have been on line. All installations are on
hog-mill waste combustors. A prototype unit was installed at
163
-------
TABLE 20- SUMMARY OF REXNORD GRAVEL BED USERS
Plant/
Date
Install.
A
6/73
B
4/75
C
1974
Gas
Capacity
Am /min
5,100
(180 M ACFM)
3,820
(135 M ACFM)
4,020
(142 M ACFM)
Gas
Temperature
°C
126
(258°F)
163
(325°F)
182
(360°F)
Pressure
Drop
cm W.C.
36
(14"W.C.)
19
(7.5"W.C.)
13
(5"W.C.)
Capital
Cost
$
750,000
2,500,000
1,400,000
Annual
Power
Cost
$
78,000
8,000
Annual
Maintenance
Cost
$
6,000
4,000
-------
i.o r
o
I—I
H
U
2:
o.i-
H
W
u
l-l
H
o;
0.01
0.1
I I
I I
I i i 1 I
I I I I I I I 1
1.0
AERODYNAMIC PARTICLE DIAMETER, ymA
10
Figure 79. Experimental grade efficiency curve of a Rexnord
gravel bed filter (McCain, 1976),
165
-------
TABLE 21. DESIGN SPECIFICATIONS OF THE SYSTEM
AS TESTED
Inlet Volume Flow: 2,266 ACM/min at 204°C
(80,000 ACFM at 400°F)
Backflush Volume Flow: 317 ACM/min at 66°C
(11,200 ACFM at 150°F)
Pressure Drop: 25.3 cm W.C. (10 in. W.C.)
Gravel Size: 4 mm (5/32 in.) x no. 6 mesh
Bed Depth: 11.4 cm (4 1/2 in.)
Bed Area: 3.72 m2/bed (40 £t2/bed)
(For a total of 59.5 m2 of bed area with
52 m2 actively filtering in normal operation)
166
-------
the Weyerhaeuser Company, Snoqualmie Falls, Washington Plant.
This unit was single down-flowing sand annulus 2.6 m (8.5 ft)
O.D. and about 1.8 m (6 ft) I.D. with an effective filtering
height of 4.9 cm (16 ft). The surface area was calculated at
34 m2 (365 ft2). The granules were angular in shape and ranged
from 3.2 mm to 6.4 mm (1/8 in. to 1/4 in.) in average diameter.
The granules moved downward through the annulus in gravity flow
at a bulk velocity of 61 to 122 cm/hr (2 to 4 ft/hr). The granule
inventory was of the order of 36 metric tons though only about
18 metric tons were in the region exposed to gas flow. The
manufacturer suggested capacity was 1,133 Am3/min (40,000 ACFM).
Hood (1976) reported the performance test data on this unit.
The experimental penetration curves are shown in Figure 80 for
three cascade impactor sampling runs. As can be seen, the unit
performed at an efficiency of from 75 to 95% in the removal of
particulates of 2 vim in diameter, at an efficiency of 70 to 90%
for particulate removal of 1 ym in diameter, and at an efficiency
of 65 to 80% for particles 0.5 ym in diameter. The efficiency of
the unit was low for the removal of particles of 0.25 ym in dia-
meter. The data indicated that in the removal of particles in
the lower size ranges the unit performed at an efficiency of
between 45 and 65%.
Under the sponsorship of the Department of Energy, Combustion
Power Company conducted parametric studies on their GBF system
to correlate the collection efficiency of the GBF with mechanical
and process parameters. Parameters studied included superficial
gas velocity, dust loading, particle size distribution, granule
diameter, granule circulation rate, bed thickness and length of
bed.
The GBF was a pilot unit and was operated at ambient tempera-
ture. Figure 81 shows the system. The granular bed material
(alumina) flows downward between two concentric cylinders. The
gas passes through the bed and is filtered by the granules. The
granules are recycled pneumatically and the collected dust par-
ticles are disengaged from the granules and sent to a conventional
baghouse.
167
-------
1.0
0.1
OS
H
tu
2
W
O-.
0.01
0.1
1 I I I
1.0
PARTICLE DIAMETER, ym
10
Figure 80
Experimental penetration curves for CPC dry
scrubber (Hood, 1976).
168
-------
GAS
IMLET
GAS
OUTLET
DISENGAGEMENT
VESSEL
FLUIDIZED BED
FLUIDIZING AIR
MEDIA RETURN PIPE
FRONT PANEL
FILTER PANEL
OUTLET PANEL
TRANSPORT PIPE
. MEDIA OUTLET PIPE
TRANSPORT AIR
EJECTOR
AIR
Figure 81. Continuous moving bed GBF.
169
-------
The performance of this device has been reported by Guillory
(1978), Wade, et al . (1978) and Wigton and Wade (1978). Test
parameters for the nominal, thick bed, thin bed, short bed, and
small collector granule configurations are listed in Table 22.
Particle concentrations ranged from 0-46 to 4.6 g/Nm3 (0.2 to
2.0 gr/SCF) . Different test dusts were used in order to vary
the mass median particle diameter from 3 to 10 ymA. The super-
ficial gas velocity was varied from 20 to 80 cm/s (40 to 160 ft/min)
The granule flow rate was varied to 0.4 to 1.6 kg of granule/kg of
air. Pressure drop ranged from 1.2 to 5.7 kPa (5 to 23 in W.C.).
Fractional efficiency curves are shown in Figures 82 through
86. The overall penetrations are correlated with the pressure
drop function and are tabulated in Table 23. The pressure drop
function was defined as:
AP C
0 = - V.
ur M
G m
where 0 = pressure drop function
AP = pressure drop, cm W.C.
UQ = superficial gas velocity, cm/s
M = media rate, kg granules/kg air
In general the CPC moving bed filter was found to be capable
of particle removal efficiencies in excess of 98% for particles
in the 1 to 10 ymA diameter range. Submicron particles were
collected at an efficiency in excess of 90% in cases with high
velocities, high loadings, and low media rates. Beds with larger
thickness to granule diameter ratios were most effective in the
capture and retention of particles in the 2 to 5 ymA diameter
range. Also, intermittent media movement was shown to improve
efficiency by a few percent. No cost data were available.
Ducon Granular Bed Filter
There is no current industrial user of the Ducon granular
bed filter. However, one test unit was installed in a refinery
170
-------
TABLE 22. TEST PARAMETERS FOR CPC MOVING BED
FILTER (from Wade, et al., 1978)
Granule Active Bed Bed Thickness
Configuration Diameter, mm Length, cm cm
Nominal 2 134.6 20.3
Thick Bed 2 134.6 40.6
Thin Bed 2 134.6 10.2
Short Bed 2 67.3 20.3
Small Medium 0.8 134.6 20.3
TABLE 23. CPC MOVING BED FILTER OVERALL
PENETRATION CORRELATION
Configuration Correlation
Nominal Ft" = 0.06 0"
Thick Bed Ft = 0.0175 0~°'7
Thin Bed Ft = 0.091 0~°'*7
Short Bed Ft" = 0.0425 0~°>S
Small Medium Ft = 0.0259 0~°'59
171
-------
tsJ
0.5
o 0.1
0.05
0.01
II1I I I I \-
I I I I I
0.5 1 5 10
AERODYNAMIC PARTICLE DIAMETER, gmA
Figure 82. Experimental grade penetration (CPC data
for nominal configuration).
1.0
0.5
0.1
0.05
0.01
TT
I I I 11
I I I I I I I I
0.5 1 5 10
AERODYNAMIC PARTICLE DIAMETER, umA
Figure 83. Experimental grade penetration (CPC data
for thick bed configurations).
-------
to control the particulate emission from the regenerator of a fluid
catalytic cracking (FCC) unit. The unit had four filter elements;
each had 14 beds. The bed was a 6.3 cm (2.5 in.) deep bed of 760
Vim sand. The face velocity was between 15 and 45 cm/s (0.5 and
1.5 ft/s). Gas temperature at the inlet was 371 to 482°C (700 to
900°F).
The overall collection efficiency reported by Kalen and
Zenz (1973) was 85 to 98% for particles with a mass median diameter
of 35 ym and a geometric standard deviation of about 4. No grade
penetration or efficiency curves were given by them.
The grade penetration curve was calculated for the Ducon
granular bed filter based on the information provided by Kalen
and Zenz (1973). Figure 87 shows the results. The shaded area
is the range of calculated points.
A high temperature and pressure design of the Ducon filter
was tested at the Exxon miniplant (Hoke, et al. (1978). The
performance data for all runs through November, 1977 are listed
in Table 24. The efficiencies are based on air inlet concentra-
tion of 2.3 g/Nm3 (1.0 gr/SCF) which is the average for the emis-
sions from the secondary cyclone. The lowest demonstrated parti-
culate outlet concentration was 68.6 mg/Nm3 (0.03 gr/SCF) however,
they were unable to maintain this level of performance for more
than a few hours of operation. At times the filtration efficiency
was very poor and the outlet particulate concentration was as
high as 700 to 1,200 mg/Nm3 (0.3 to 0.5 gr/SCF). Fractional
efficiency data are presented in Figure 88.
EVALUATION
Granular bed filters perform in many respects in a manner
analogous to fiber filtration systems. The major difference
appears to be the size differences between the fiber used in
fabric filters and the granules in granular bed filters.
175
-------
1.0
0.05
o
3
w
0.1
0.05
0.01
0.5 1
5 10
PARTICLE DIAMETER, u
i i i
15
Figure 87 . Fractional penetration curve for Ducon
granular bed.
1.0
0.5
0.1
0.05
0.01
0.1
Figure
I I I I I I I I I
I I I I I I I
UG = 45 cm/s
dc = 0.04 cm
Z = 3.8 cm
I I I I I I I II
A I I I 1 J
0.5 1
PARTICLE DIAMETER,
10
Fractional penetration for Ducon GBF (from Bertrand, et
al., 1977).
-------
TABLE 24. GRANULAR BED FILTER PERFORMANCE
(FROM BERTRAND, ET AL., 1977)
Outlet Concentration
Run Number
54
57
59 (Sample 1)
59 (Sample 2)
59 (Sample 3)
61
62.1
62.3
63 (Sample 1)
63 (Sample 2)
63 (Sample 3)
64 (Sample 1)
64 (Sample 2)
64 (Sample 3)
65 (Sample 1)
65 (Sample 2)
66
gr/SCF
0.69
0.04-0.08
0.08
0.28
0.54
0.46
0.03
0.21
0.05
0.07
0.12
0.28
0.29
0.27
0.05
0.06
0.06
g/m3
1.57
0.09-0.18
0.18
0.64
1.23
1.05
0.07
0.48
0.11
0.16
0.27
0.64
0.66
0.61
0.11
0.14
0.14
Collection Efficiency*
31.0
92.0-96.0
92.0
72.0
46.0
54.0
97.0
79.0
95.0
93.0
88.0
72.0
71.0
73.0
95.0
94.0
94.0
* Based on a 2.3 g/Nm3(1.0 gr/SCF) inlet concentration
177
-------
Generally, for low to medium temperature applications, granu-
lar bed filters are not economically competitive with fabric fil-
ters even though the gas flow capacity of granular bed filters
is much higher than that of fabric filters. Because granular
bed filters are not mass produced and the weight of the bed re-
quires special support, the cost of granular bed filters is much
higher than that of fabric filters. Also, granular bed filters
generally have lower efficiencies than fabric filters.
The best potential applications of granular beds as devices
to control particle emissions are in situations which require
the control of effluents under both corrosive and high tempera-
ture conditions. GBFs are most successful in controlling emissions
of particulates that agglomerate easily as in the case of cement
dust.
Cleaning Methods
During the operation of a granular bed filter, d.ust deposits
in the interstices of the bed and on the surface. It is necessary
to clean the dust from the bed to prevent it from saturating
the bed and causing high pressure drop.
Depending on bed structure, different cleaning methods are
used. Fixed bed GBFs, such as Ducon GBF and Rexnord GBF,
apply a reverse gas flow to blow the dust out. In the continu-
ously moving bed and intermittently moving bed GBFs, the bed
is cleaned by removing both the granules and the dust from the
filter and repacking the filter with clean granules.
The advantages and disadvantages of present GBF systems
are discussed in the following paragraphs.
Fixed Bed GBFs
The fixed bed GBFs are usually horizontal beds. Beds are
cleaned by reverse gas flow. The reverse flushing gas may be
used to fluidize the bed and elutriate the dust deposits as
in the case of the Ducon granular bed system.
The advantage of fixed bed GBFs is that they require
no granule recirculation system and therefore have a lower opera-
ting cost.
178
-------
The disadvantages are:
1. Leakage or bypassing;
2. Plugging of retaining grids, and loss of granules
when retaining grids are not used;
3. Efficiency decreases with time (particle seepage); and
4. Ineffective bed cleaning.
Leakage or Bypassing -
Dust loosened in the cleaning cycle may be carried directly
to the stack when the bed is put back on line. This problem has
been described by Kalen and Zenz (1973) with the Ducon GBF and
by McCain (1976) with the Rexnord GBF.
Plugging of Retaining Grids -
The collection efficiency of the granular bed can be in-
creased by using finer grades of granules. However, grids with
openings smaller than the granules need to be used. This in-
creases the tendency to accumulate dust deposits on the grid
and eventually causes plugging. This problem was experienced by
Rexnord GBF users and by Exxon Research and Engineering Company
with the Ducon granular bed installed at their pressurized fluid-
ized bed combustor miniplant. At Exxon, the inlet retaining grid
of the Ducon granular bed filter plugged quickly even when the
inlet grid was a 10 mesh screen. The deposited dust on the screen
could not be blown out by reverse gas flow. Shutdown of the plant
was required to remove the dust manually.
In solving this plugging problem, Exxon has eliminated the
inlet retaining grid by increasing the free board above the bed
(Bertrand, et al., 1977). However, this arrangement causes granules
to be blown out and lost from the filter beds. In subsequent
tests Exxon lowered the blow back gas velocity. Some indications
of particle buildup in the filter beds at the low blow back velo-
cities were noted. In one run about 35% of the bed was fly ash
which was not blown out of the bed. It was also found that if
the filter beds were overloaded with fine particles, the particles
tended to move to the top of the filter bed during blow back where
they prilled, forming larger spheres which could not be removed
179
-------
except at much higher blow back velocities. Exxon has tried using
heavier granules, however, some of the material was still lost
during blow back.
Efficiency Decreases with Time -
For non-agglomerating dust, the blow back cleaning technique
will decrease the collection efficiency of the bed with each
cleaning cycle. During blow back, collected dust redisperses
into the gas stream. However, the fine particles will remain in
suspension. This will increase the particle loading of fine par-
ticles in the gas stream. Since the blow back gas is cleaned by
other beds and the collection efficiency of the bed is less than
100%, the fine particle penetration will progressively increase
with each cleaning cycle, (i.e., the efficiency of the bed goes
down). Exxon (Bertrand, et al. 1977) has noticed this decrease
in efficiency with time. In one test by Exxon, the filtration
efficiency was fairly high at the start of the run. The outlet
particulate loading was about 0.12 g/m3 (0.05 g/SCF) . The outlet
loading increased somewhat during the run to 0.17 g/m3 (0.07 g/SCF)
and finally to 0.25 g/m3 (0.1 g/SCF) at the end of the 12 hour test
Ineffective Bed Cleaning -
Cleaning by reverse gas flow is not an effective cleaning
method especially when the cleaning duration is short. The clean-
ing efficiency is even lower when particles are charged.
Generally, an adhesive force exists between the collected
dust and the granules. To separate the collected dust from
granules by reverse gas flow, the aerodynamic drag force must be
higher than the adhesive force between the dust and granules. For
those particles which have a strong adhesive force, as in the case
of charged particles, dust will not be removed completely from the
granules by the reverse gas flow alone. Thus, dust will gradually
accumulate in the bed. In testing the Ducon GBF, .Exxon (Bertrand
et al., 1977) discovered a large accumulation of fly ash in the
bed, uniformly distributed through the filter medium.
180
-------
The accumulation of dust in the filter bed has some signi-
ficant consequences. The dust can approach the outlet section of
the filter and plug the retaining screens. Once the dust reaches
the retaining screen, it could be entrained and blown out of the
filter bed during the filtration step.
Continuously Moving Bed GBFs
This method is normally limited to vertical panel filters.
Dust and granules are continuously removed at the bottom of the
GBF. The advantage of this type of bed structure is that the
collected dust and granules are separated outside the GBF. This
method does not increase the particle loading in the gas stream
and the cleaning is more effective compared to reverse gas flow.
The disadvantages are:
1. Costly solid handling system
2. Solids distribution
3. Solid flow
4. Particle reentrainment
5. Erosion
6. Heat loss
Solid Handling -
A costly granule circulation system and dust/granule separa-
tion system is required. Under low to medium gas temperature
environments, collected dust and granules may be separated by
shaking. Cleaned granules are transported to the top of the
panel by mechanical means. In high temperature and pressure
(HTP) applications, these methods are not feasible. Other
methods of transport and separation are required.
Under HTP conditions, granules could be transported by
pneumatic means or some improved mechanical means. Depending
on the mass flow rate ratio, pneumatic transport may require a
large quantity of compressed gas. This will increase the capi-
tal cost and operating cost.
Dust and granules under HTP conditions are difficult to
separate by mechanical shaking because of the requirement for a
suitable metal bellows or seal for the moving mechanism. It
181
-------
may be possible to separate dust and granules by sending them
through a baffled "rattler." Dust would be shaken off the granules
when the granules hit the baffle. The dust would then be elutriated
out of the rattler with air and later cleaned by means of a conven-
tional separator such as a baghouse and cyclone at lower tempera-
ture. This setup requires the investment of low temperature
secondary cleanup systems.
Solid Flow -
The downward movement of the solids can create a dead zone
near the filter surfaces because some granules are retained by
the louvers. These zones could eventually be saturated with dust
and lead to plugging of the louvers.
Recently, Combustion Power Company made a design change to
alleviate this problem. In the new design, gas flows from the
inside core radially outward through the bed. The inlet louvers
are replaced with slotted panels. During operation, solids flow
downward as well as spill through the slots. This design eliminates
the dead zone.
Particle Reentrainment -
The grinding of the granules due to the relative motion of
the filter granules can dislodge the collected particles and
allow them to be reentrained into the gas stream. The reentrain-
ment rate depends on the granule recirculation rate and filtration
gas velocity. At high recirculation rate, collected particles
are easier to dislodge. At too low a recirculation rate, the bed
may be saturated with collected dust. Depending on the inlet
particle loading, there exists an optimal granule recirculation
rate. Under the sponsorship of the Department of Energy, CPC
conducted a parametric test (Wade, 1977, Guillory, 1977) on their
GBF. The granule recirculation rate was one of the parameters
CPC studied. They showed that low recirculation rate and inter-
mittent media movement improved collection efficiency by a few
percent. This improvement could be attributed to a lower particle
reentrainment.
182
-------
Solids Distribution -
Solids distribution to the filter panels may present some
difficulties. It is not easy to distribute the solids evenly
to the panels.
Erosion -
Solids retaining elements will be subjected to erosion by
the moving granular bed. Selection of materials that can resist
erosion in HTP conditions becomes a problem. Under the sponsor-
ship of ERDA, Combustion Power Company designed and constructed
a full scale filter to treat the total gas flow from the CPU-400
fluidized bed combustor at 1,000 kg/min (2,200 Ib/min). The gas
temperature was at 704°C (1,300°F). The retaining louvers were
made of RS 330 steel. During a shakedown test in December 1975,
the louvers had a structural failure in about twenty hours of
operation.
Heat Loss-
Granules are withdrawn from the bed for cleaning. To keep
the granules hot, substantial energy may be required for HTP
conditions. If the granules are recirculated by pneumatic means
it is necessary to pre-heat the transport air to minimize heat
loss from granules.
Intermittently Moving Bed
Intermittent movement is normally limited to vertical panel
filters. The granules are intermittently removed in a cross-
flow arrangement, as in CCNY's panel bed. The advantages of this
type of bed structure are external granule/dust separation and
minimum disturbance to the rooting cake. A rooting cake is the
foundation on which the surface cake is formed. The surface cake
is formed readily without disturbing the rooting cake and filtra-
tion efficiency is higher.
The disadvantages are the same as discussed earlier for the
continuously moving bed. In addition, the intermittently moving
bed suffers the following disadvantages.
1. Low gas capacity
2. High operating cost and heat loss
183
-------
Low Gas Capacity -
During cleaning, about two to three layers of granules are
removed from the bed. To prevent the dust from being carried
deep into the bed by the gas, the filtration velocity should be
kept as low as possible to reduce the aerodynamic drag force.
CCNY usually operates the panel bed filter at about 15 cm/s
(30 ft/min). This velocity is about one third the velocity used
in the fixed bed and continuously moving bed GBFs. Thus, more
filtration area is required.
High Operating Cost and Heat Loss -
To prevent the dust from penetrating deep into the bed, sur-
face layers are ejected frequently. Depending on dust loading,
the surface layer has been removed as frequently as every 30
seconds in the operation of the CCNY panel bed GBF. At 30
second "puff-back" frequency, the volume of the reverse gas flow
is about 1% of the total volume of gas treated. The reverse
gas flow has a pressure of about 140 kPa (20 psi) higher than the
pressure of the gas to be treated. The costs to operate the com-
pressor which supplies the reverse gas flow may be substantial
if the gas is to be treated at high pressure.
This cleaning method has another drawback. The "puff-back"
gas will have to be pre-heated in order to prevent significant
heat loss. To maintain the gas temperature by pre-heating the
puff-back gas, the energy requirement is high.
184
-------
SECTION 8
POTENTIAL FOR HTP APPLICATIONS
Fluidized bed coal combustion and low-BTU coal gasification
are among the advanced energy processes which require high tempera-
ture and high pressure (HTP) particulate cleanup. In addition, HTP
cleanup might be required for other high temperature and/or high
pressure processes as reported by Parker and Calvert (1977).
The suitability of granular bed filters (GBF) for controlling
particulate emissions from advanced energy processes is not limited
by the gas temperature and pressure. By properly selecting the
adequate granules and structural materials, the granular bed filters
should be capable of operating at any temperatures and pressures
encountered in advanced energy processes.
CLEANUP REQUIREMENTS
In order to evaluate the potential of GBFs for HTP cleanup,
it is necessary to consider the following:
1. The cleanup requirements.
20 The performance characteristics of GBFs.
30 The particulate size distribution and concentration in the
inlet gas.
4o Capital and operating cost estimates for GBFs.
5o Permissible costs for HTP gas cleanup.
The points listed above are discussed in this section, A
summary and conclusions section is given at the end of this section.
Parker and Calvert (1977) have reviewed and evaluated the
cleanup requirements for various processes. The conditions for HTP
particle collection are summarized in Table 25. As can be seen
from Table 25 gas temperatures range up to 1,100°C (2,000 F) and
pressures range up to 70 atm. Granules, such as quartz sand and
ceramics', can handle these extreme conditions.
The Ducon GBF, Combustion Power Company moving bed filter,
and the CCNY panel bed filter, all can be designed to be operated
185
-------
TABLE 25. CONDITIONS FOR HIGH TEMPERATURE AND PRESSURE PARTICULATE COLLECTION
PROCESS
TEMPERATURE
°C
PRESSURE
atm
TYPICAL GAS COMPOSITION
mol %
EXPECTED
PARTICULATE
COMPOSITION
00
Open cycle
coal-fired
gas turbine
Fluidized
bed coal
combustion
Coal gasifi-
cation
02 blown
Air blown
FCC regener-
ator
Metallurgical
furnaces
MHD power
generation
650-1,000
800-900
150-1,100
300-800
250-1,000
300-800
4-10
-1-20
-1-70
-1-3
~1
831 N2, 15% C02 , 2% 02, H20,
SOx, NOX, CO, and gaseous
hydrocarbons
80% N2, 10% C02> 6% 02,
4% H20, + S02, NO, CO
30% H2, 25% CO, 15% C02,
20% H20, 3% CH,,, H2S, N2
50% N2, 12% H2, 20% CO,
10% H20, 6% C02, + CHu,
H2S
68% N2, 5% CO, 31 02,
8% C02, 16% H20, + NOX,
SOX, NH3, HCN, aldehydes,
hydrocarbons
N2, C02, 02
coal ash, unburnt
carbon
60 wt % ash, 30%
unburnt carbon,
10% sorbent
ash, unburnt
carbon, sorbent,
possibly tar
catalyst dust
depends on cata-
lyst type, commonly
silica and alumina
very fine metal
fume
K2COs seed par-
ticles
-------
at HTP. The Rexnord GBF can be operated at high temperature
but not at high pressure. Rexnord does not recommend that their
GBF be used at HTP. They have designed a new type of filter for
these conditions. The new design is similar to the CPC moving
bed GBF.
The use of granular bed filters for HTP applications is
limited by the particulate and gaseous pollutant removal effi-
ciencies. Particulate cleanup requirements for HTP processes
vary depending on the intended use of the gas. If it is to be
vented, the gas must be cleaned sufficiently to meet the emission
standards. Current new source performance standards are 43 mg/MJ
(0.1 lb/106 BTU), however, a stricter standard of 13 mg/MJ (0.03
lb/106 BTU) has been proposed.
If the hot gas is to be expanded through a gas turbine, then
the gas must meet the turbine requirement for cleanliness. A gas
containing dust particles can severely erode and corrode turbine
blades and other internal components. Also, deposition of dust
particles on the turbine blades can impair the aerodynamic
performance of the turbine.
A large number of research investigations have been reported
which deal with turbine blade erosion and deposition problems.
Much of this work was done in connection with military gas tur-
bines for helicopter and tracked-ground vehicle engines. Similar
research has also been conducted with industrial gas turbines.
Generally, it is believed that large particles (over 2-5 ym
diameter) cause severe erosion damage and must be removed. Par-
ticles smaller than 1-2 ym diameter cause much less erosion
damage. However, there is a scarcity of data concerning the tol-
erance of turbines for fine particles.
From the available data on turbine tolerances for particulate
matter, it appears that effectively all particles larger than
about 2 ym must be removed from the gas. It has been suggested
(Westinghouse, 1974) that a mass loading of 370 mg/Nm3 (0.15
gr/SCF) for particles smaller than 2 ym would allow a satisfactory
turbine life. The particulate removal requirements imposed by
187
-------
the gas turbine limitations may not be as stringent as the emission
standard. If there were a sufficient loading of particulate
smaller than 2 urn, it would be possible for the gas to be cleaned
sufficiently to protect the turbine while still exceeding the
emissions regulation of 0.1 lb/106 BTU (approximately 0.05 gr/SCF).
Recently Sverdrup and Archer (1977) proposed that to protect
the turbine, the particulate concentration should be no more than
5 mg/Nm3 (0.002 gr/SCF) and there should be no particles larger
than 6 ym in diameter.
PREDICTED GBF PERFORMANCE
Figure 89 gives the predicted granular bed performance
for several conditions. It shows the effects of temperature and
pressure on particle penetration through a granular bed filter.
The predictions were made for a bed packed with 400 urn diameter
granules to a depth of 3.8 cm (1.5 in.)- Other assumptions were:
an approaching gas velocity of 45 cm/s, no surface cake
formed, the collected dust was uniformly distributed in the
bed, the average void fraction of the bed between cleaning cycles
was 0.25, and the particle density was 1.5 g/cm3.
As concluded in a report by Calvert and Parker (1977), high
temperature and pressure particle collection is more difficult
than at low temperatures. The predicted penetration for the gran-
ular bed filter at 875°C and 10 atm is much higher than that for
ambient conditions.
If the particle size distribution of a source were known, it
would be possible to predict whether a granular bed filter would
be able to meet various cleanup requirements. The equation relat-
ing the fractional penetration (for a specific particle diameter)
to the overall penetration is:
Pt = Ptd £(dp)d(dp)
188
-------
1.0
z
o
PL,
•X
O
I—I
H
W
2
W
OH
1-J
u
I— I
H
OS
<
P-
0. 5
0.1
0.05
0.01
0.1
GRANULE DIAMETER: 400v»m
BED DEPTH: 3.8 cm
= 45 cm/s
= 1.5 g/cm3
= 0.25
10 a tin
870°C
1 atm
870°C
iiii
0.5 1.0
PARTICLE DIAMETER, ym
I I 1
Figure 89. Predicted GBF performance
10
189
-------
where Ft = overall penetration, percent or fraction
Pt, = penetration for particles with diameter, d , fraction
f(d ) = particle size frequency distribution
d = particle diameter, ym or cm
Particle Size and Concentration
Information on particle size distribution and mass loading
of particles leaving HTP sources is scarce. The best available
data are those reported by Exxon Research and Engineering Company
(Hoke, 1976, 1977). Exxon has a pressurized fluidized bed coal
combustor miniplant. The gas leaves the combustor at a temperature
of about 870°C (1,600°F) and a pressure of about 10 atm0 First it
passes through a primary cyclone which removes larger particles
(including unburnt carbon) and recycles these particles to the
combustor. The gas leaves the primary cyclone and passes through
a secondary cyclone„ This removes more large particles and reduces
the mass loading to the order of 2.5 g/Nm3 (1 gr/SCF)„ A Ducon
GBF is connected to the outlet of the secondary cyclone to further
clean the gas stream,
Hoke (1976) reported the particle size distribution at the
secondary cyclone exit. It was obtained by sieve and Coulter
counter analysis. The particles have a mass mean diameter of 8 ym
and geometric standard deviation of 2.7 (Figure 90).
Overall collection efficiency was calculated graphically from
equation (97)for this size distribution and for a granular bed
packed to a depth of 3,8 cm with 400 ym diameter granules„ The
approach velocity of the gas was assumed to be 45 cm/s0 The over-
all penetration was calculated to be 208% (collection efficiency
97.2%). Therefore, the predicted emissions will be 0.07 g/Nm3
(0002& gr/SCF) and it is in compliance with the current standard
for particulate emissions (001 lb/106 BTU or about 1014 g/Nm3)0
However, it will not meet the proposed new standard (0,03 lb/106
BTU or about 0.04 g/Nm3).
190
-------
50
w
H
W
o
w
H
OS
10
1.0
0.5
JULY 1976
DATA
MAY 1977
DATA
10 20 30 40 50 60 70
WEIGHT PERCENT UNDERSIZE, %
80
90
Figure 90
Particle size distributions from Exxon miniplant
fluidized bed coal combustors. From monthly
report #77, July 1976 and report #86, May 1977.
191
-------
As revealed in Figure 89 , the predicted penetration decreases
rapidly around 1 ym. The predicted penetration for 2 ym particles
is effectively zero. Therefore, there should be no particles big-
ger than 2 ym in the GBF exhaust. The predicted loading of 0.07
g/Nm3 (0.028 gr/SCF) meets the turbine requirement suggested by
Westinghouse (1974); but, it does not meet the requirement proposed
by Sverdrup and Archer (1977).
Hoke (1977) reported some new data on particle size distribu-
tion at the secondary cyclone exit. The mass median diameter is
3.5 ym and the geometric standard deviation is 2.9.
The overall penetration for this size distribution will be
4.99 (95.1% overall collection efficiency). The emission will be
0.12 g/Nm3 (0.049 gr/SCF). This still satisfies the current emis-
sion standard (approximately 0.06 gr/SCF) and also meets the tur-
bine requirement suggested by Westinghouse).
From the above discussion, it appears that improvements in
GBF designs are required in order to improve collection efficiency
and resolve operation problems. However, theoretical collection
efficiencies appear to be sufficient to meet the current emissions
regulations for particulates. Performance may be satifactory for
protecting gas turbines, however, this will depend strongly on
the amount of submicron particles a turbine can tolerate.
GASEOUS POLLUTANTS
In pressurized fluidized bed combustion, besides the emission
of particulates, gaseous pollutants such as S02 and alkali metal
compounds may also be present in the flue gas from the combustor.
The presence of S02 might not have any ill effect on the operation
of the gas turbine«, The S02 may be removed from the gas stream
at HTP with a GBF with dolomite or alumina as granular material
or be removed by conventional means after the gas has been expanded
through the turbine.
The presence of alkali metal compounds, such as sodium and
potassium chlorides will contribute to hot corrosion of the gas
turbine. Therefore, to protect the turbine, the concentrations
192
-------
of alkali metal compounds in the flue gas have to be reduced to a
tolerable level. A direct method for accomplishing this is by
controlling the combustion to minimize the evolution of alkali
metal compounds. An alternative method is to use GBFs with some
sort of sorbents as granular material to remove particulates and
alakli metal compounds simultaneously.
Limited work has been done in this area. Swift et al. (1977)
have shown that activated bauxite, which is a thermally treated
high alumina content natural bauxite ore, was efficient at removing
NaCl vapor at 900°C and atmospheric pressure.
As with the case of particulate matter, the acceptable level
of alkali metal vapor in flue gas stream is not well established.
Therefore, it is impossible at this time to speculate whether the
GBFs are able to remove harmful vapors to acceptable levels.
PRELIMINARY COST ESTIMATE
Granular bed filter technology is still in the developmental
stage and there are no proven designs to be used in HTP cleanup.
Many of the cost factors have not been established. Therefore,
it is extremely difficult to do a detailed cost analysis.
In the following section, the relative capital and operating
costs between a HTP fixed bed GBF, a HTP intermittently moving
bed, and a HTP moving bed system are presented. The cost esti-
mates were performed for the GBF layouts presented in Figures 72,
81 and 91. It should be noted that these designs might not be
the optimum.
Basis
In this report, the estimate was based on the cleanup re-
quirement for one gas turbine. The seventeen-stage axial flow
compressor is designed to develop a pressure ratio of approximately
10 to 1 while using 345 kg/s (761 Ib/s) air at IOS (International
Organization for Standardization) ambient conditions (298.33°K
dry bulb temperature with 60% relative humidity). The four-stage
expander is designed to operate with a turbine inlet temperature
of 1,233°K (1,760°F), and under these conditions the net output
193
-------
has been calculated to be approximately 71 MW (Beecher, et al.
1976). If the turbine inlet temperature were 1,143°K instead of
1,233°K and the turbine efficiency remained the same, the turbine
power output would be about 66 MW. The flow rate of 760 Ib/s at
ambient conditions is approximately 6,700 m3/min at 870°C and 10
atm (237,000 ACFM at 1,600°F and 10 atm) and 7,220 m3/min at 960°C
and 10 atm (255,000 ACFM at 1,760°F and 10 atm).
The following were used as the basis for the cost estimate:
1. Inlet gas temperature: 870°C (1,600°F)
2. Inlet gas pressure: 10 atm
3. Outlet gas pressure: 1 atm
4. Net turbine power output: 66 MW
5. Gas volume flow rate: 6,700 Am3/min at 870°C
6. Total power output for combined cycle plant: 355 MW
Capital Costs
Since the GBF's are non-standard fabricated units, the esti-
mated fabricated cost was based on the costs of raw materials
and labor requirements. Costs for design, administration, con-
tingency, engineering, etc. were not included. Auxiliary equip-
ment was priced based on vendor's quotations. Equipment cost was
bare module cost, not installed cost. The cost for this study was
developed in terms of fourth quarter 1978 U.S. dollars.
Fixed Bed GBF -
The fixed bed GBF is designed for a superficial gas velocity
of 40 cm/s (80 ft/min). Total bed area required is 275 m2 (2,960
ft2). Eight GBFs are required. Each GBF has 48 filter elements
and each filter element has 16 beds. Total bed area per GBF is
34.3m2(370 ft2).
The GBF shell is a pressure vessel. The dimensions are shown
in Figure 91 . It is made of 3.8 cm (1.5 in.) thick carbon steel
lined with 15 cm (6 in.) thick refractory. There is also an inner
of 0.32 cm (1/8 in.) thick type-316 S.S.
The filter element consists of two perforated concentric tubes
made from type-316 S.S. Beds of granules are stacked vertically
194
-------
BACKWASH MANIFOLD
AND VALVES
FILTER ELEMENT
Figure 91. Fixed bed GBF.
195
-------
within the annulus. The O.D. of the outer tube is 30 cm (1 ft)
and the I.D. of the inner tube if 19 cm (7.5 in.).
Material costs used in the estimates are: carbon steel
$1.00/kg ($0.45/lb); 316 S.S. $9.90/kg ($4.48/lb); castable re-
fractory $0.74/kg ($0.34/lb). Fabrication labor costs are:
carbon steel, $1.00/kg ($0.45/lb); 316 S.S., $3.96/kg ($1.80/lb);
castable refractory, $0.74/kg ($0.34/lb).
The fabricated costs of the GBF system are as follows:
Pressure Vessel $960,000
Filter Elements 693,000
Blowback Valves, Piping § Lock Hopper 925,000
Ductwork (refractory lined) 247,000
Compressor 179,000
Total $3,004,000
The estimated GBF bare module cost is $448.36/Am3 ($12.68/AGFM).
The compressor indicated is for the supply of reverse gas
flow during cleaning. It has a capacity of about 102 m3/min
(3,600 SCFM) and the developed pressure is 1,388 kPa (200 psig).
Continuously Moving Bed GBF -
The arrangement of the moving bed is shown in Figure 81.
The bed is a single downflowing annulus 2.6 m (8.5 ft) O.D. and
2.3 m (7.5 ft) I.D. The height of the filtering surface is 4.9
m (16 ft). The bed is placed in a 3.7 m (12 ft) diameter pressure
vessel.
Each GBF can handle a gas capacity of 1,130 Am3/min (40,000
ACFM) at a superficial gas velocity of 50 cm/s (100 ft/s). There-
fore, six GBFs are required to treat the gas for one gas turbine.
The pressure vessel consists of three layers. The outside
shell is 3.8 cm thick carbon steel. The inner layer is 16 gauge
(1/8") 316 S.S. In between there is a 15 cm (6 in.) thick layer
of refractory. The material and labor costs to fabricate the
pressure vessel are the same as presented in the last section.
The outer panel of the granular bed is a perforated wall and
the inner retainer is a slotted panel. Since granules are
196
-------
continuously moving downward, the bed retaining walls should be
made of erosion resisting materials; e.g., Hastelloy C, RA 330, and
RA 333. The material cost used in this cost estimate is $39.40/kg
($17.90/lb). Fabrication labor cost is estimated to be $7.92/kg
($3.60/lb).
Under HTP conditions, it is impractical to separate dust and
granules by shaking apparatus and to transport granules by mechani-
cal means. Therefore, pneumatic transport and solid/dust separa-
tion techniques were designed into the GBF system. It is assumed
that the weight of granules recirculated is equal to the weight
of gas treated; i.e., granule recirculation rate is 345 kg/s
(760 Ib/s) . Granules are withdrawn from the bed by means of an
ejector and are transported to a granule holding tank above the
GBF. The mass flow rate ratio of granules to transport air is
assumed to be 20 g/g.
The granules flow from the granule holding tank by gravity
into a baffled rattler section. In the rattler section, the dust
is knocked off the granules and elutriated by an air stream which
is also used as the fluidizing air to help the flow of granules.
The mass flow rate ratio of granules to the fluidization air is
assumed to be 80 g/g. All granule transport lines are carbon
steel pipe lined with refractory and erosion resisting metal
liners.
The fabricated cost of the GBF system is as follows:
Pressure Vessel $ 791,000
Filter Retaining Walls § Stiffeners 2,156,000
Hot Gas Piping 52,000
Granule Cleaning Ductwork 110,000
Ductwork 185,000
Compressor 784,000
Secondary Collector (incl. ductwork) 150,000
Total $4,228,000
The estimated fabricated cost is about $631.04/Am3 ($17.84/
ACFM), which is 1411 higher than that for the fixed bed system.
The cost of continuously moving bed systems can be reduced
if a better method for granule recirculation is found. In the
197
-------
estimate, low bulk density pneumatic transport is used for granule
transport. Other transport methods such as dense-phase pneumatic
transport may be more cost efficient.
Some testing is required to determine whether the dense-phase
vertical pneumatic transport of 1-2 mm solids is feasible. If it
is feasible, the capital cost (compressor cost) and operating
cost will be lower than predicted in this estimate.
Intermittently Moving Bed GBF -
One possible arrangement for the intermittently moving bed
GBF is shown in Figure 72. Each square filtration module is 0.56
m x 0.56 m x 4.6 m (22 in. x 22 in. x 15 ft). Active filtration
area per module is 5.57 m (60 ft , each side is 1 ft x 15 ft).
Filtration velocity is about 15 cm/s (30 ft/min). Therefore,
total filtration area required is 21,970 m2 (236,500 ft2) and the
number of filtration module required is 132. The filter elements
are made of Hastelloy "C". Since the panel is a louvered wall,
material requirement is 8 times the unit length of the panel
(Wu, 1977).
It is possible to place 17 modules in a pressure vessel 4.88
m (16 ft) in diameter. Thus, eight GBFs are required for each gas
turbine.
As in the case of the continuously moving bed GBF, granule
circulation and dust/granule separation are accomplished by pneu-
matic means.
The cost data used in the estimate were the same as used
previously. The estimated costs are as follows:
Pressure Vessel $1,050,000
Filter Elements 5,186,000
Ductwork 250,000
Blowback Cleaning System 132,000
Compressor '92,000
Lock Hopper § Granule Cleaning System 720,000
Total $7,430,000
198
-------
The estimated fabricated cost is about $1,108.96/Am3 ($31.35/
ACFM) which is 247% higher than that of the fixed bed system and
about 176% higher than that of the continuously moving bed system.
The greatest cost item is that of the filter element. The
cost can be reduced if the GBF is run at a higher filtration velo-
city. The filtration velocity used in this estimate is that
recommended by Squires (Wu, 1977). Research work is required
to determine whether it can be run at higher filtration velocities.
Operating Power Costs
The operating power costs for the GBF systems described
above were estimated in this study and include those due to the
following:
1. Pressure drop across the bed.
2. Power requirement to operate the compressor which is used
either for cleaning or for pneumatic transport of solids.
3. Cost to heat the transport air and cleaning air to the
bed temperature.
4. Heat loss to the surroundings.
Pressure drop and heat losses cause the gas temperature and
pressure at the turbine inlet to be lower than it would be with-
out cleanup. Lower gas temperatures and pressures reduce the tur-
bine power output which in turn lowers the revenue.
The fixed bed operating power cost (not including deprecia-
tion) is the lowest among the three systems. The operating power
costs of the continuously moving bed and the intermittently moving
bed are about 7.4 times and 4 times higher than the fixed bed,
respectively.
More than 60% of the moving bed operating cost is due to the
compressor for pneumatic transport. The moving bed operating
cost can be lowered if better solids transportation methods are
found.
In the intermittently moving bed system, the high operating
cost is also due to the compressor. However, the compressor is
not for solids transport as in the continuously moving bed system.
It is used for "puff-back" air. The "puff-back" air requirement
199
-------
per puff is 0.28 m3 (10 ft3) for each square filter unit. "Puff-
back" frequency depends on dust loading. It is usually about one
puff every half minute.
PERMISSIBLE COSTS
High temperature and pressure (HTP) particulate control is
an important factor in determining the economic feasibility of
many advanced energy processes. The energy process which relies
most strongly on HTP gas cleanup is the pressurized fluidized bed
(PFB) combustion process. The PFB process was described in detail
as part of the Energy Conversion Alternatives Study (EGAS) - a
cooperative effort of the Energy Research and Development Adminis-
tration (now the Department of Energy), the National Science
Foundation and the National Aeronautics and Space Administration.
For the PFB process to be economically competitive it is
necessary to recover energy from the HTP effluent gas by expanding
it through a gas turbine. To protect the turbine and not lose
energy, it is essential that the gas be cleaned with a minimum
loss of temperature and pressure.
The General Electric and Westinghouse Phase II EGAS reports
(Brown et al. 1976 and Beecher et al. 1976) present detailed
designs and cost estimates for PFB boiler power plants. The re-
sults are summarized in Table 26 (from Lewis Research Center, 1977)
The overall cost of electricity for the General Electric design
is presented in Figure 92 as a comparison with other advanced
processes considered in EGAS Phase II. The reference steam cycle
is a conventional coal-fired boiler power plant with wet lime
stack gas scrubbers (Brown, 1976) .
The advanced steam PFB designs assumed Ducon granular bed
filters would be suitable for HTP particle collection. The EGAS
reports were not explicit as to what percentage of the cost of
electricity is attributable to the granular bed filters, although
about 20% of the capital investment was for the cleanup system.
Figure 92 shows that a cost difference of 5 to 6 mills/kWh exists
between the PFB with cleanup and the reference steam cycle. We
200
-------
TABLE 26. SUMMARY OF EGAS PHASE II PERFORMANCE AND COST RESULTS
System and contractor
1 - AFB/steam
(General Electric)
2 - PFB/ steam
(General Electric)
3 - PFB/ steam
(Westinghouae)
4 - PFB/poLaeeium/eteam
(General Electric)
5 - AFB/closed-cycle gas
tu rbine/or ganic
(General Electric)
6 - Low-Btu gasifier/gas
turbine/ steam
(General Electric)
7 - Low-Btu gaslfler/gas
turbine/eteam
(Westinghouse)
8 - Semiclean-fuel-fired
gas turbine/steam
(Westinghouse)
9 - Scmiclean- fuel- fired
gas turbine/eteam
(General Electric)
10 - Coal/MHD/ steam
(General Electric)
11 - Low-Btu gaBifler/moltcn-
carbonate fuel cell/stoam
(United Technologies Corp.)
Net
power,
MW
814
904
679
996
476
585
786
874
847
1932
635
Efficiency, percent
Thermo-
dynamlc
43.9
41.3
42.3
47.8
50.1
44.2
48.5
53.6
52.7
54.0
53.6
Power-
plant
35.8
39.2
39.0
44.4
39.9
39.6
46.8
52.2
51.1
49.8
49.6
Over-
all
35.8
39.2
39.0
44.4
39.9
39.6
46.8
38.6
37.8
48.3
49.6
Capital cost,
$/kWe
EGAS
ground
rules
632
723
549
934
1232
771
614
329
418
720
593
Constant
mid- 1975
dollars
447
411
401
660
899
562
448
256
306
478
433
Cost of electricity,
mills/kW-hr
ECAS
ground
rules
31.7
34.1
28.1
39.9
49.3
35.1
29.1
26.0
29.5
31.8
28.9
Constant
mid- 197 5
dollars
25.8
27.4
23.5
31.2
38.6
28.6
23.9
23.7
25.9
24.1
23.9
201
-------
Reference Steam Cycle
COST OF ELECTRICITY (Mills/kWhl
10 20 30 40
<-
50
Advanced Steam - AFB1814.3 MWe)
Advanced Steam - PFB (903.8 MWe)
Closed Gas Turbine:
Organic Bottoming - AFB (474.1 MWe)
Potassium Topping Cycle - PFB 1993.8 MWe>
Open Cycle Gas Turbine Combined:
Water Cooled - Semiclean Fuel (846.5 MWe)
Open Cycle Gas Turbine Combined;
Air Cooled - LBTU Gasifier (584.8 MWe)
Own Cycle MHO (mi 2 MWe)
Key (Contribution to Total):
Capital I I Fuel
V////A Maintenance and Operating I
Figure 92. Summary G.E. Comparison of Cycles (Cost of
Electricity at an assumed capacity factor
of 65%).
202
-------
estimate that the capital cost for cleanup would be about 8.5
mills/kWh so that cost difference between the PFB without cleanup
and the reference steam plant with cleanup is about 14 mills/kWh.
Therefore, one can conservatively state that a cost on the order
of a few mills/kWh would make the PFB process economically competi-
tive for further development. There is also the advantage of fuel
conservation which is due to the higher overall efficiency of the
PFB process as compared to a conventional power plant.
SUMMARY AND CONCLUSIONS
Particulate matter removal from gas at HTP is desirable for
any process which involves the use of a turbine on the HTP gas
stream. The degree of particulate cleanup required to protect
the turbine is uncertain. Published estimates of turbine require-
ments range from a maximum 370 mg/Nm3 (0.15 gr/SCF) for particles
smaller than 2 ym diameter (Westinghouse, 1974) to a maximum
loading of 5 mg/Nm3 (0.002 gr/SCF) and no particles larger than
5 ym in diameter (Sverdrup and Archer, 1977).
Particle data obtained on the Exxon PFBC Miniplant range
from 2.5 g/Nm3 of fly ash with d = 9 ym and ag = 2.7 to the
same loading with d = 3.5 ym and a = 2.9. The A.P.T. perfor-
mance model when applied to the larger particle size distribution
and the Ducon GBF in use on the Miniplant predicts outlet loadings
around 70 mg/Nm3 and no particles larger than 2 ym diameter,
which meets the Westinghouse turbine requirements and present
New Source Performance Standard (NSPS) (about 140 mg/Nm3) but
does not meet the Sverdrup and Archer turbine requirements or
the proposed NSPS (abour 40 mg/Nm3).
For the smaller particle size distribution the A.P.T. model
predicts that the Ducon GBF would give an outlet loading of
about 120 mg/Nm3, which is close to the present NSPS.
A criterion for the permissible cost of HTP particulate
removal can be estimated from the performance and cost estimate
data available from the EGAS studies (Figure 92 ) . The PFBC
plant is a good example because its feasibility may depend on a
203
-------
satisfactory HTP cleanup system. The power generating difference
between a PFBC and a conventional steam power plant is about
14 mills/kWh. A gas cleanup cost on the order of a few mills/kWh
should be acceptable and even higher costs may be justifiable.
Cost estimates were made for three GBF systems for HTP clean-
up of the combustion gas in a 355 MW combined cycle power plant.
The gas turbine provides 66 MW of the plant capacity. Of the three
GBF systems, the fixed bed system has the lowest estimated capital
and operating costs. The capital costs of the continuously moving
bed system and the intermittently moving bed system are predicted,
respectively, 141% and 2471 higher than the fixed bed system.
With regard to the operating power cost, the continuously moving
bed system and the intermittently moving bed system are about 7.4
times and 4 times higher than that of the fixed bed system.
By assuming a cost factor of 4 in the estimation of costs of
engineering, installation, site preparation, contractor's fee,
contingency, construction, and working capital cost, the capital
investment of a continuously moving bed system would be $16,912,000;
or $47.64/kWh of plant capacity. The operating power cost of the
continuously moving bed system is about 0.48 mills/kWh of plant
capacity. For a 10 year life, depreciation charge is about 0.55
mills/kWh of plant capacity. Thus, the estimated total operating
cost would be about 1 mill/kWh for the continuously moving bed
system. This cost figure is economically competitive when com-
pared to the EGAS Phase II results.
However, it should be noted that no granular bed filter sys-
tems have been demonstrated to have high enough collection effi-
ciency and reliable performance to satisfy HTP gas cleanup re-
quirements. More development and redesign are necessary. Con-
sequently, any estimates of cost and process economics should be
considered highly speculative at this time. Adequate comparisons
between GBF system economics must be based on proven pilot plant
designs.
204
-------
SECTION 9
RECOMMENDATIONS FOR FUTURE RESEARCH AND DEVELOPMENT
The primary objective of evaluating the feasibility of granu-
lar bed filters (GBF) for the collection of particulates at high
temperature and pressure (HTP) has been achieved in this study.
It has been shown that GBFs are capable of operating at HTP.
Whether the GBF meets the cleanup requirements depends on the
application. Present GBFs have the potential to meet the current
emission standards. However, unless aided by other collection
mechanisms, the present designs are not likely to meet the proposed
NSPS for boilers or the turbine requirements proposed by Sverdrup
and Archer (1977). More research and development work is required
to improve the performance of GBFs. Future research and develop-
ment work is needed in the following areas:
1. Efficiency improvements.
2. Bed cleaning methods
3. Other potential problems, such as sintering of the
granules, granule transport method, etc.
EFFICIENCY IMPROVEMENT
The present investigation has shown that existing GBFs may
not have a high enough collection efficiency for fine particles,
especially when operating at high temperature. There are a few
studies reported in the literature which show that the collection
efficiency of the bed may be increased by:
1. Electrostatic augmentation
2. Cake filtration
205
-------
Electrostatic Augmentation
The filtration efficiency can be enhanced by electro-
static augmentation. If the filtration medium is immersed
in an electrostatic field, the dust particles will be driven
in a direction that tends to increase the probability of
impact between particles and the filter medium. Limited
work has been done with electrostatically augmented granular
beds. A.P.T. has carried out experimental work to study the
effect of charging the bed and/or particles on the collection
efficiency. This work will be presented in a separate
report.
These small-scale experiments have demonstrated that the
GBF collection efficiency can be improved greatly by charging
the bed or by charging both the bed and particles. It is
possible for the GBF to achieve a 99+% collection efficiency
when an electrostatic field is imposed on the bed.
For industrial applications, the capital cost of electro-
statically augmented GBFs will be higher than without augmenta-
tion. However, the increase in capital outlay is compensated
by the reduction in operating cost. The electrostatically
augmented GBF can be operated at much lower pressure drop by
using a shallow bed and the energy cost to maintain the electro-
static field is small (5 x 10~6 mills/kWhe).
The next steps in evaluating the electrostatically augmented
GBF should move toward a pilot-scale test on an actual source,
such as., a fluidized bed coal combustor. More laboratory-scale
studies are required to refine the technology and establish
design criteria for the pilot plant. Areas which need further
research include:
1. bed charging method.
2. electrode configurations.
3. bed structure.
4. HTP electrode insulation.
5. effects of polarity on efficiency.
6. cleaning methods
206
-------
Most of these effects can be evaluated in the laboratory
and the most promising combination can be tested on a pilot-
scale plant.
Cake Filtration
There is good reason to believe that the GBF with a good
surface cake would have a much higher filtration efficiency
than the original clean bed. This would result from increased
particle collection by sieving.
There is no published information on the efficiency of cake
filtration. A small-scale laboratory research program could
generate the needed information on cake filtration. One approach
could be to build a cake on the surface of the GBF and then
measure the collection efficiency using monodisperse particles.
The effects of cake thickness, pressure drop, particle size
distribution of the cake, and superficial gas velocity would be
studied.
Further research is needed to determine the GBF operating
conditions which result in cake formation. Whether or not there
will be a surface cake depends largely on whether a rooting cake
will be formed. During a filtration cycle, particles will de-
posit in the interstices of the bed and will coat the granules.
If the granules of the bed are large, the coatings or dendrites
will not bridge to create the rooting cake. On the other hand,
if the granules are sufficiently small, dendrites will bridge
to form an internal cake.
The formation of surface cake has been observed at CCNY
(Wu, 1977) for ambient temperatures. However, at high tem-
peratures, there was no indication that a surface cake was
formed. Exxon Research and Engineering Company (Bertrand,
et al., 1977) reported the same experience. No surface cake
was observed in the Ducon GBF in their HTP fluidized bed coal
combustor miniplant.
207
-------
Westinghouse Research Laboratories reported some collection
efficiency data for their laboratory scale GBF under ambient
conditions (Ciliberti, 1977). The GBF utilized fluidized air
to clean the bed. The reported overall collection efficiencies
were high and the grade efficiency curves calculated from cas-
cade impactor data were flat (i.e., almost independent of particle
diameter). These phenomena indicated that a surface cake was
present.
Surface cake formation at high temperature has been observed
for filtration on a ceramic cloth medium (Shackleton and Kennedy,
1977). Therefore, the formation of surface cake appears to be
dependent upon the substrate, although it is influenced somewhat
by temperature. Research work is required to study the cake
formation process at high temperatures. We recommend a pilot-
scale study using actual combustion flyash. Variables to be
studied should include bed granule diameter, bed structure, gas
temperature and pressure, and bed operating conditions.
A small-scale laboratory study on cake filtration and the
pilot plant study on the cake formation process could be studied
concurrently.
BED CLEANING METHODS
To prevent the bed from becoming saturated with collected
dust, it is necessary to clean the bed either periddically or
continuously. The present bed cleaning methods were evaluated
in this report. Major research and development work is needed
to improve the efficiency of the present cleaning methods and to
decrease the operating costs.
Present bed cleaning methods share a common problem - des-
truction of rooting cake. In the continuously moving bed system,
the relative motion of bed granules prevents the formation of
the internal cake. In the reverse gas flow cleaning method,
the bed is either fluidized or stirred. Thus internal cakes
208
-------
are destroyed or deteriorated. The rooting cake needs to be
rebuilt before a surface cake can be formed.
Since the rooting cake is the foundation of surface cake
and the formation of a surface cake enhances the collection effi
ciency of the GBF, it is desirable to preserve the rooting cake
during the cleaning cycle. Research work is required to develop
cleaning methods that remove the surface cake and preserve the
rooting cake.
OTHER RESEARCH RECOMMENDATIONS
There are many operational problems and uncertainties which
need to be resolved before HTP granular bed filters can be
considered sufficiently reliable and economical for commercial
application. These problems include:
1. How to prevent particle seepage through the bed (during
cleaning or filtration).
2. How to reduce temperature losses (especially during
cleaning).
3. How to improve the efficiency and reduce the cost of
granule regeneration and recirculation.
4. How to reduce pressure drop across bed.
5. How to prevent attrition of granules causing particle
reentrainment.
Resolving these problems will not only help solve the
HTP particle collection problem, but will improve granular
bed filter technology for many other applications, especially
where hot, corrosive gases are encountered.
RECOMMENDED RESEARCH PROGRAM
Of the above mentioned research and development recommen-
dations, electrostatic augmentation is presently the most
proven approach to improve the performance of the GBFs. It has
209
-------
been shown that the electrostatically augmented GBF has the
potential to meet the most stringent cleanup requirements under
ambient conditions. It is expected that the same statement will
hold for HTP conditions where higher electric field can be imposed
on the beds. Therefore, electrostatic augmentation is recommended
as the approach having the best probability of success, based on
present knowledge.
Another potential advantage of the electrostatic GBF is that
it achieves high collection efficiency without the benefit of cake
filtration. In actual applications, there might be filter cakes
and further research may show the way to promoting surface cake.
The presence of surface and rooting cakes will increase the col-
lection efficiency of the GBF even more.
The capital cost of an electrostatic GBF will be slightly
higher than that of regular GBFs. The additional cost is due to
the requirement of a high voltage power supply and HTP insulation
for the electrodes. However, the operating cost could be lower
due to a lower drop in pressure. The estimated operating cost
for the Ducon GBF is about 31.8 mills/m3/min (0.9 mills/SCFM).
For an electrostatically augmented Ducon, GBF, the operating cost
would be 27.9 mills/mVmin (0.79 mills/SCFM) to maintain the same
efficiency.
A detailed program to demonstrate the feasibility of using
electrostatic augmentation to improve GBFs for particulate con-
trol at high temperature is described below. We recommend a
study of the electrostatically augmented GBF on a pilot plant
scale of about 14.2 Am3/min (500 ACFM). To duplicate actual
industrial applications, fresh test dust should be produced in-
stead of regenerated dust. Since GBFs will be used in advanced
energy processes, it is desirable to test the electrostatically
augmented GBF on these processes. A good approach would be to
use an atmospheric fluidized bed combustor.
The GBF should be designed in such a way that it is easy to
change from one configuration to- another. Bed cleaning can be
achieved either by fluidization or by continuously withdrawing
granules and dust from the bed.
210
-------
To aid in the design of the pilot plant, some small-scale
experimental work should be conducted concurrently,
In outline, the objectives consist of the following tasks
1. Conduct small-scale experiments to obtain design
information.
2. Design the pilot plant.
3. Fabricate, install and start up the pilot plant.
4. Prepare a detailed test plan describing:
. a. The proposed test matrix,
b. The measurement techniques to be used.
c. The data handling methods.
5. Conduct test programs.
6. Analyze data, conduct engineering and cost analyses
of various configurations.
7. Based on the above analyses, design and estimate the
cost of a GBF system for HTP applications.
8. Recommend a test program to demonstrate a full-scale
GBF system on an HTP source.
211
-------
REFERENCES
Anderson, D.M. and L. Silverman, "Mechanisms in Electrostatic
Filtration of Aerosols with Fixed and Fluidized Granules,"
report by the Air Cleaning Laboratory, Harvard University
to the U.S. Atomic Energy Commission, No. NYU-4615, 1958.
See also, "Development of a Triboelectrified Fluidized Bed
for Aerosol Filtration," 5th Atomic Energy Commission Air
Cleaning Conference, p. 140, USAEC, TID-7551, 1957.
Avco, Inc. "Evaluation of Granular Bed Devices," Avco Applied
Technology Division, AVATD-0107-69-RR, June 1969.
Balasubramanian, M. and A. Meisen, "A Note on the Diffusional
Deposition of Aerosol Particles in Packed Bed," J. of Aerosol
Science, £: 461-463 (1975).
Beecher, et al., "Energy Conversion Alternatives Study (EGAS)
Westinghouse Phase II Final Report, Volume III - Summary and
Advanced Steam Plant with Pressurized Fluidized Bed Boilers,"
NSF/RA-760590, NTIS PB 268-558, October 1976.
Berry, F.J. arid J. Fournier, German Patent #676508 (1939).
Bertrand, et al., "A Regnerative Limestone Process for Fluidized
Bed Coal Combustion and Desulfurization," Monthly Report
no. 87, 1977.
Billings, C.E. and J. Wilder, "Handbook of Fabric Filter Tech-
nology, Volume I," NTIS PB 200-648 (1970).
Bohm, L. and S. Jordan, "On the Filtration of Sodium Oxide
Aerosols by Multilayer Sand Bed Filters," J. of Aerosol
Science, I: 311-318 (1976).
Borguardi, R.H., R.E. Harrington, and P.W. Sparte, "Filtration
Characteristics of Fly Ash," J. of APCA, 18: 387 (1968).
Brown, et al., "Energy Conversion Alternatives Study (EGAS)
General Electric Phase II Final Report, Volume II - Advanced
Energy Conversion Systems, Conceptual Designs," NASA-CR
134949, NTIS PB 268-247, December 1976.
Brown, D.H., "Conceptual Design and Assessment of a Utility
Steam Plant with Conventional Furnace and Wet Lime Stack
Gas Scrubbers," NASA CR-134950, December 1976.
212
-------
Calvert, S. and R. Parker, "Effects of Temperature and Pressure
Calvert, S., Chapter 46 in "Air Pollution," A.C. Stern, editor,
3rd edition, Academic Press, New York, 1968.
Carney, S.D., U.S. Patent #2343401 (1944).
Ciliberti, D.F., Westinghouse Research Laboratories, personal
communication.
Gorman, J.C., and G.R. Fox, "ECAS-General Electric Phase II
Final Report, Volume I: Executive Summary," NASA-CR 134949,
NTIS #PB 268-455, December 1976.
Dennis, R. , et al., "Special Incineration Studies - Institutional
Design," 6th AEC Air Cleaning Conference, July 7-9, 1959.
TID-7593, pg. 344-364 (1960).
Ergun, S., "Fluid Flow Through Packed Columns," Chemical Engineering
Progress, 4_£: 89-94 (1952).
Fiechter, L.B., British Patent #135216 (1919).
Fiechter, L.B., British Patent #163039 (1922).
Figueroa, A.R., "Aerosol Filtration by Fixed and Fluidized
Granular Beds," Ph.D. Dissertation, University of Cincinnati,
(1975).
Figueroa, A.R., and W. Licht, "Filtration of Submicron Particles
by Fixed and Fluidized Granular Beds," Paper #33b presented
at the 81st AIChE National Meeting, Kansis City, Missouri,
April 11-14, 1976.
Fournier, J., British Patent #450048 (1936).
Friedlander, S.K., "Mass and Heat Transfer to Single Spheres
and Cylinders at Low Reynolds Numbers," AIChE Journal,
3: 43-48 (1957).
Fuchs, N. and A. Kirsch, "The^ Effect of Condensation of a Vapor
on the Grains and of Evaporation from Their Surface on the
Deposition of Aerosol in Granular Bed," Chem, Eng. Science,
'20: 181-185 (1965).
Gebhart, J., C. Roth, and W. Stahlhofen, "Filtration Properties
of Glass Bead Media for Aerosol Particles in the 0.1-2 ym
Size Range," J. of Aerosol Science, 4_: 355-371 (1973).
Goldman, L., "Experiences in the Use of Pebble-Bed Filters in
Inorganic Chemical Plants," Staub, 2£: 449 (1964).
213
-------
Goren, S., in "Granular Bed Filter Development Program: Monthly
Report for August 1977" by K.E. Phillips, prepared for the
U.S. Energy Research and Development Administration under
contract #EF-77-C-01-2579, October 1977.
Guillory, J.L, "Filtration Performance of a Moving Bed Granular
Filter: Experimental Cold Flow Data," paper presented at
the 5th International Conference on Fluidized Bed Combustion,
Washington, D.C., December 12-14, 1977.
Guillory, J.L, "Granular Bed Filter Development Program,"
prepared for the U.S. Dept. of Energy, FE-2579-15, 1978.
Hoke, R.C., et al., "A Regenerative Limestone Process for
Fluidized Bed Coal Combustion and Desulfurization," Monthly
Report no. 77, July 1976.
Hood, K.T., "Evaluation of the Combustion Power Company Moving
Gravel Bed Dry Scrubber on the Control of Particulate
Emissions from a Hog Fired Boiler," NCASI special report,
September 1976.
Jackson, S. and S. Calvert, "Entrained Particle Collection in
Packed Beds," AIChE Journal, L2: 1075-1078 (1966).
Kalen, B. and F.A. Zenz, "Filtering Effluent from a Cat Cracker,"
Chem. Eng. Progress, 66: 67-71 (1973).
Klarding, N., British Patent #153263 (1921).
Knapsack-Griesham Aktiengesellschaft, Koln, Germany, British
Patent #851776 (1960).
Knettig, P. and J.M. Beeckmans, "Capture of Monodispersed Aerosol
Particles in a Fixed and in a Fluidized Bed," Canadian
J. of Chemical Engineering, 52: 703-706 (1974).
Kovach, I.L., and D.G. Hanan, "Combined Dust, Aerosol and Vapor
Removal Efficiency of Packed Activated Carbon Beds,"
Staub, 3_0: (12), 35-39 (1970).
Lapple, C.E., "Stack Contamination - 200 Areas, Interim Report,"
HDC-611, August 5-October 12, 1948.
Lee, K.C., "Filtration of Redispersed Power-Station Fly Ash by
a Panel Bed Filter with Puffback," Ph.D. Dissertation,
The City University of New York, (1975).
Lee, K.C., et al., "Panel Bed Filter," EPRI-AF-560, 1977.
214
-------
Levenspiel, 0., "Chemical Reaction Engineering," Wiley 8 Sons,
New York, 1972.
Leith, D., S.N. Rudnick, and M.W. First, "High Velocity, High
Efficiency Aerosol Filtration," NTIS PB 249-457, 1976.
Leith, D. and M. First, "Performance of a Pulse-Jet Filter at
High Filtration Velocity, I - Particle Collection," J.
of APCA 27_: 534-539 (1977).
Levi, J.P. and J. Blume, French Patent #763077 (1934).
Lewis Research Center, "Evaluation of Phase II Conceptual De-
signs and Implementation Assessment Resulting from the
Energy Conversion Alternatives Study (EGAS)," NASA TMX-
73515, NTIS PB 270-017, April 1977.
Lynch, G.E., U.S. Patent #1766221 (1930).
Lynch, G.E., Fuel Economist, 12: 47 (October 1936).
Marple, V., "The Fundamental Study of Inertial Impactors,"
Ph.D. Thesis, University of Minnesota, 1970.
McCain, J.D., "Evaluation of Rexnord Gravel Bed Filter," EPA
600/2-76-164, NTIS PB 225-095, June 1976.
McFee, D.R., and J. Sedlet, "Plutonium-Uranium-Molybdenum Fume
Characteristics and Sand Filtration," J. of Nuclear Energy,
22_i 641-650 (1968).
Mercer, T.T. and R.G. Stafford, "Impaction from Round Jets,"
Ann. Occupational Hygiene, 12: 41-48 (1969).
Mercier, E. and M. Ehlinger, U.S. Patent #2493356 (1950).
Miyamoto, S. and H. Bonn, "Filtration of Airborne Particulates
by Gravel Filters: I. Initial Collection Efficiency of a
Gravel Layer, " J. of APCA, 20; 1051-1054 (1974).
Miyamoto, S. and H. Bohn, "Filtration of Airborne Particulates
by Gravel Filters: II. Collection Efficiency and Pressure
Drop in Filtering Fume," J. of APCA, 25_: 40-43 (1975).
Nordstrom, 0., Swedish Patent #56098 (1922).
Nordstrom, 0., British Patent #217113 (1924).
Orr, C., Jr., "Particulate Technology, " p. 420, MacMillan Com-
pany, New York (1966).
215
-------
Parker, R. and S. Calvert, "High Temperature and High Pressure
Particulate Control Requirements," EPA 600/7-77-071, July
1977.
Payatakes, A.C., "Model of Aerosol Particle Deposition in Fibrous
Media with Dendrite-Like Pattern: Application to Pure Inter-
ception During Period of Unhindered Growth," Filtration
and Separation, 15; 602 (1976a).
Payatakes, A.C., "Model of the Dynamic Behavior of a Fibrous
Filter; Application to Case of Pure Interception During
Period of Unhindered Growth," Powder Technology, 14;
267 (1976b).
Payatakes, A.C., "Model of Transient Aerosol Particle Deposition
in Fibrous Media with Dendrite Pattern," AIChE Journal,
23^: 192-202 (1977).
Payatakes, A.C. and C. Tien, "Particle Deposition in Fibrous Media
with Dendrite-Like Pattern: A Preliminary Model, " J. of
Aerosol Science, 1_: 85-100 (1976).
Paretsky, L., L. Theodore, R. Pfeffer, and A.M. Squires, "Panel
Bed Filters for Simultaneous Removal of Fly Ash and Sulfur
Dioxide: II. Filtration of Diluted Aerosol by Sand Beds,"
J. of APCA, 21^: 204-209 (1971).
Paretsky, L.C., "Filtration of Aerosols by Granular Beds,"
Ph.D. Thesis, City University of New York (1972).
Perry, J.H. (editor)/'Chemical Engineer's Handbook,"McGraw-
Hill, New York, 1973.
Phillips, K.E. "Granular Bed Filter Development Program: Monthly
Report for August 1977," prepared for the U.S. Energy
Research and Development Administration under contract
no. EF-77-C-01-2579, October 1977.
Ranz, W.E. and J.B. Wong, "Impaction of Dust and Smoke Particles,"
Ind. Eng. Chem. 44: 1371-1381 (1952).
Research Cottrell, U.S. Patent No. 2,990,912, July 4, 1961.
Schmidt, E.W., J. Gieseke, P. Gelfand, T.W. Lugar, and D.A. Furlong,
"Filtration Theory for Granular Bed," J.. of APCA, 28: 143-146,
1978. —
Shackleton, M. and J. Kennedy, "Ceramic Fabric Filtration at
High Temperatures and Pressure," paper presented at the
EPA/ERDA Symposium on High Temperature/Pressure Particulate
Control, Washington, D.C., September 20-21, 1977.
216
-------
Sharapov, K.A., V.V. Leonov, I.L. Saklarnova, A.F. Skvartsov,
and N.G. Braginets, "Study of a Multigradient Electro-
magnetic Filter for the Dry Cleaning of Gases," Stal'
963-964 (1975), Ref. Zh. Khim. GI775 (1976).
Solvay, E., British Patent #18573 (1889).
Squires, A.M. and R.A. Graff, "Panel Bed Filters for Simultaneous
Removal of Fly Ash and Sulfur Dioxide: III. Reaction of
Sulfur Dioxide with Half-Calcined Dolomite," J. of APCA,
21.: 272-276 (1971).
Squires, A.M. and R. Pfeffer, "Panel Bed Filters for Simultaneous
Removal of Fly Ash and S07: I. Introduction," J. of APCA,
2Q_: 523 (1970). L
Stephan, D.G., G.W. Walsh, and R.A. Herrick, "Concepts in Fabric
Air Filtration," AIHA Journal, 2±: L (1960).
Stern, A.C., H.W. Zeller, and A.I. Schekman, "Collection Effi-
ciency of Jet Impactors at Reduced Pressures," Ind. and
Eng. Fundamentals, !_: 273 (1962).
Strauss, W. , and M.W. Thring, "Studies in High Temperature Gas
Collecting," J. of the Iron and Steel Institute, 196:
62 (1960).
Sverdrup, E.F. and D.H. Archer, "The Tolerance of Large Gas
Turbines to Rocks, Dusts, and Chemical Corrodants," pre-
sented at EPA/ERDA Symposium on High Temperature and Pres-
sure Particulate Control, Washington, D.C. September 1977.
Swift, M.W.. et al. "Plans and Studies on Flue Gas Cleaning
and Particulate Monitoring in PFBC," Proceedings of the
EPA/DOE Symposium on High Temperature/ High Pressure Parti-
culate Control, EPA 600/9-78-004, September 1977.
Taub, R. , "Filtration Phenomena in a Packed Bed Filter," Ph.D.
Dissertation, Carnegie-Mellon University, 1970.
Thomas, J.W. and R.E. Yoder, "Aerosol Size for Maximum Penetra-
tion Through Fiberglass and Sand Filters," Archives Ind.
Health, 121: 545 (1956a) .
Thomas, J.W. and R.E. Yoder, "Aerosol Penetration Through a Leak
Shot Column, A Method of Particle Size Evaluation," AMA
Archives Ind. Health, 13^: 550 (1956b) .
Thomson, T. and N. Nisbet, British Patent #216675 (1924).
Thring, M.W. and W. Strauss, "The Effect of High Temperatures
on Particle Collection Mechanisms," Trans. Instru. Chem.
Eng. 41: 248-253 (1963).
217
-------
Tien, C., C.S. Wang and T.D. Barot, "Chain-Like Formation of
Particle Deposition in Fluid Particle Separation," Science,
196: 983-985, 1977.
Veron, M., U.S. Patent No. 2,564,316 (1951).
Wade, G.L., "Performance and Modeling of Moving Granular Bed
Filter," paper presented at EPA/ERDA Symposium on High
Temperature/Pressure Particulate Control, Washington, D.Ci,
September, 1977.
Wade, G., H. Wigton, J. Guillory, G. Goldbach, and K. Phillips,
"Granular Bed Filter Development," Final report prepared
for U.S. Dept. of Energy, No. FE-2579-19, 1978.
Wang, C.S., M. Beizail, and C. Tien, "New Concepts of Particle
Deposition from Suspensions Flowing Past a Collector,"
paper #18b, presented at the AIChE 83rd national meeting
Houston, Texas, March 1977.
Westinghouse Electric Corp., "Clean Power Generation from
Coal," O.C.R., 84, NTIS: PB 234-188, April 1974.
Wigton, H. and G. Wade, "Granular Bed Filter Development Program,"
report prepared for the U.S. Dept. of Energy, No. FE-2579-
18 (1978).
Wilson, E.J. and C.J. Geankoplis, Ind. Engineering Chemistry,
5: 9 (1966).
Wu, M.S., Air Pollution Technology, Inc., Personal Communication,
1977.
Zahradnik, R.L, J. Anyigbo, R.A. Steinberg, and H.L Toor,
"Simultaneous Removal of Fly Ash and S02 from Gas Streams
by a Shaft Filter Sorber," Env. Science and Tech., 4: 663-
667 (1970).
218
-------
/», TECHNICAL REPORT DATA
(Please read Instructions on the reverse before completing)
1. REPORT NO.
EPA-600/7-79-020
2.
3. RECIPIENT'S ACCESSION-NO.
4. TITLE AND SUBTITLE
Evaluation of Granular Bed Filters for High-temper-
ature/High-pressure Particulate Control
5. REPORT DATE
January 1979
6. PERFORMING ORGANIZATION CODE
7. AUTHOR(S)
Shui-Chow Yung, Ronald Patterson, Richard Parker,
and Seymour Calvert
8. PERFORMING ORGANIZATION REPORT NO
9. PERFORMING ORGANIZATION NAME AND ADDRESS
Air Pollution Technology, Inc.
4901 Morena Boulevard, Suite 402
San Diego, California 92117
10. PROGRAM ELEMENT NO.
1NE624
11. CONTRACT/GRANT NO.
68-02-2183
12. SPONSORING AGENCY NAME AND ADDRESS
EPA, Office of Research and Development
Industrial Environmental Research Laboratory
Research Triangle Park, NC 27711
13. TYPE OF REPORT AND PERIOD COVERED
Final; 9/76 - 9/78
14. SPONSORING AGENCY CODE
EPA/600/13
15. SUPPLEMENTARY NOTES JERL.RTp projeCt OfHC6r IS
2925.
Drehmel, MD-61, 919/541-
is. ABSTRACT
repOrt gives resuitg of a critical review and evaluation of the status
of granular bed filter (GBF) technology. GBFs can operate at the high temperatures
and pressures normally encountered in advanced energy processes. In theory, the
filters can clean the gas to meet the current New Source Performance Standard if
certain operating difficulties are overcome. Use of a combined- cycle power plant
with GBFs for particulate cleanup is economically competitive when compared to a
conventional power plant. However, successful and reliable performance at high tem-
peratures and pressures has not yet been demonstrated. None of the design equations
reported in the literature are adequate to predict the collection efficiency of a GBF
at face velocities normally encountered in the field. A performance model was
developed in this study to predict collection efficiency for GBFs. A small scale
experiment was performed to obtain additional performance data for GBFs.
17.
KEY WORDS AND DOCUMENT ANALYSIS
DESCRIPTORS
b.lDENTIFIERS/OPEN ENDED TERMS
c. COS AT I Field/Group
Pollution Dust
Gas Scrubbing Aerosols
Gas Filters Electric Power
Granular Materials Plants
High-temperature Tests
High-pressure Tests
Pollution Control
Stationary Sources
Granular Bed Filters
Particulate
13B
07A,13H
13K
11G
14B
11G
07D
10B
18. DISTRIBUTION STATEMENT
Unlimited
19. SECURITY CLASS (ThisReport)
Unclassified
20. SECURITY CLASS (Thispage)
Unclassified
!1. NO. OF PAGES
231
22. PRICE
EPA Form 2220-1 (9-73)
219
------- |