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TRANSCRIPT
THE EFFECT OF SOLID RESIDENCE TIME ON
BIOMASS GASIFICATION YIELDS
by
RONG-CHI WANG, B.S. in Ch.E.
A THESIS
IN
CHEMICAL ENGINEERING
Submitted to the Graduate Faculty of Texas Tech University in
Partial Fulfillment of the Requirements for
the Degree of
MASTER OF SCIENCE
IN
CHEMICAL ENGINEERING
Approved
Accepted
August 1981
-w
ACKNOWLEDGMENTS
The author expresses his deep appreciation to Dr. S. R. Beck,
Dr. L. D. Clements and Dr. R. W. Tock for their support and technical
advice during the course of this work.
Also, he would like to thank Jein-Fong Yu, Steve Kromer, Steve
Duran, Randy Cotton, Jay Scott and others who helped with this
project.
The financial support of the United States Department of Energy
(Contract No. DE-AC04-79ET20041) is gratefully acknowledged.
n
V
TABLE OF CONTENTS
PAGE
ACKNOWLEDGMENTS ii
LIST OF TABLES v
LIST OF FIGURES. . vi
CHAPTER 1 INTRODUCTION 1
CHAPTER 2 LITERATURE REVIEW 3
CHAPTER 3 EQUIPMENT AND PROCEDURE 21
Experimental Apparatus 21
Feed Section 21
Reactor Section 21
Tar Collection Section 27
Downstream Section 29
Operational Procedure 29
Analytical Procedure 31
Feedstock 32
CHAPTER 4 DISCUSSION OF RESULTS 33
Material Balances 33
Temperature Profiles 42
Particle Size Distribution 42
Determination of Mean Solid Residence Time 44
Carbon Conversion 49
Product Gas Yield and Compositions 57
Effect of Other Variables 66
CHAPTER 5 CONCLUSIONS AND RECOMMENDATIONS 70
LITERATURE CITED 72
• • •
/
PAGE
APPENDIX A 74
APPENDIX B 76
APPENDIX C 77
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Table
Table
Table
Table
Table
Table
Table
Table
Table
Table
Table
Table
Table
Table
Table
Table
Table
Table
Table
1
2
3
4
5
6
7
8
9
10
11
12
13
14
15
16
17
18
19
LIST OF TABLES
PAGE
Important Reactions in Gasification 11
Parameters Used to Calculate Gas-Phase
Residence Time 14
Operating Conditions 34
Elemental Analysis of Solids and Gas 36
Waste Water Analysis from Kao 37
Adjusted Material Balances 38
Hydrogen Balances 39
Oxygen Balances 40
Carbon Balances 41
Mean Particle Size and Particle Size Uniformity for Feed and Char + Cyclone Fines on Run 123,
124, 125, 126, 128 48
Mean Solid Residence Time 50
Average Solid Residence Time on Run 123 51
Average Solid Residence Time on Run 124 52
Average Solid Residence Time on Run 125 53
Average Solid Residence Time on Run 126 54
Average Solid Residence Time on Run 128 55
Gas Yield and Compositions 58
Data for Analyzing Gas Yield, H2/CO Mole Ratio and Carbon Conversion with SAS 67 Errors in Correlations 68
LIST OF FIGURES
PAGE
Figure 1 Gasification Processes and Their Products 4
Figure 2 Effect of Reactor Temperature on H2/CO
Mole Ratio 6
Figure 3 Effect of Oxygen on Gas Composition , 7
Figure 4 Total Product Gas Yield from Sawdust 8
Figure 5 Carbon Conversion as a Function of Reactor Temperature . 9
Figure 6 Major Processes Occurring in an Updraft Gasifier 10
Figure 7 Gas Production vs. Gas Residence Time for Various Gas Phase Temperature, (a) 500°C, (A) 600° (O) 650°, (•) 700°, (X) 750°C 15
Figure 8 A Model of Fluidized Bed, Particles with
Wide Size Distribution 18
Figure 9 SGFM Pilot Plant 22
Figure 10 Dimensions of Feed Hopper and Screw Feeder 23
Figure 11 Reactor Section 24
Figure 12 Reactor Dimensions 25
Figure 13 Tar Collection Section 28
Figure 14 Downstream Section 30 Figure 15 Temperature Profiles of Run 123, 124, 125,
126, 128, 131 43
Figure 16 Particle Size to Cumulative Wt. % Under Size for Run 123 45
Figure 17 Particle Size to Cumulative Wt. % Under Size for Runs 124-126 46
Figure 18 Particle Size to Cumulative Wt. % Under Size for Run 128 47
VI
I (
PAGE
Figure 19 Effect of MSRT on Carbon Conversion 56
Figure 20 Effect of MSRT on Product Gas Yield 59
Figure 21 Effect of MSRT on H2, CO, CO2 Yield 60
Figure 22 Effect of MSRT on CH4, C2H4 Yield 61
Figure 23 Effect of MSRT on C2H5, C2H2 Yield 62
Figure 24 Effect of MSRT on H2/CO Mole Ratio 63
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CHAPTER 1
INTRODUCTION
In many processes involving gas-solid reactions, fresh particles
are continually fed to a fluidized bed. The reacted particles are
either discharged through an overflow pipe or entrained by the gases.
If we wish to control and predict the behavior of a fluidized bed on a
commercial scale unit, it is M^T'J important to know the residence time
distribution of the solids in the reactor.
The SGFM (Synthesis Gas From Manure) reactor is a fluidized bed in
which biomass is gasified. The char that accumulates in the SGFM reac
tor can be removed through a center port opening in the bottom distri
bution plate. An air-operated ram, with a variable ratio of time up to
time down, is used to prevent any bridging of the char in the discharge
line and control the rate of char removal. The previous work done with
the reactor to develop a data base for the gasification of wood resi
dues did not consider the cycle time of the air-operated ram as a vari
able. We think that varying the ram cycle time should have an influence
on the solids hold-up In the reactor, which is the factor to control the
solid residence time in the bed.
For a gas-solid reaction, the length of particle hold-up, or resi
dence time, in the bed determines the extent of reaction of the solids
with the gas, even though this varies from particle to particle. For
example, some particles are discharged immediately after feeding while
others are discharged only after remaining in the bed for a long time.
Hence, for the prediction of extent of chemical reaction, it becomes
imperative to know the distribution of residence time of the particles
in the reactor and the mean solid residence time.
The objective of this study is to determine the effect of solid
residence time on gas yield, gas compositions and carbon conversion in
the SGFM pilot plant.
CHAPTER 2
LITERATURE REVIEW
In the development of synthetic fuels, much attention has centered
on biomass as a source material for the necessary synthesis gas or
chemicals. Gasification of biomass has been undertaken in a number of
ways as shown in Figure 1. The SGFM pilot plant at Texas Tech Univer
sity operates as an air or oxygen gasifier. It was designed and built
in 1975. Initially, production of ammonia synthesis from cattle manure
was studied. Cattle manure was chosen to help solve environmental
problems. This study indicated that there was also the potential for
the production of ethylene. Ethylene is important because it is the
largest volume petrochemical in the world (1). Using data from the
SGFM pilot plant, a conceptual design and economic analysis of ammonia
production from manure was performed by Bechtel National Inc. (2).
These results were promising and justified further work.
The SGFM pilot plant was used to evaluate other biomass feedstocks
such as oak sawdust, corn stover and cotton gin trash (3). This study
demonstrated the SGFM process could convert several biomass materials
to a medium calorific value gas in an air-blown gasifier. Reactor temp
erature was the primary controlling variable in determining gas yield,
but the air-to-feed ratio was also significant. Beck and Wang (4) re
ported that gas yields from oak sawdust were higher than for cattle
manure and confirmed that gasification of dry sawdust was an exothermic
reaction.
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Production of methanol synthesis gas from oak sawdust by oxygen-
blown gasification was investigated by Yu in the SGFM gasifier (5).
The results of Yu's study showed that both the average reactor tempera
ture and the oxygen-to-daf (dry, ash-free) feed ratio affected the pro
duct gas yield. The H2/CO mole ratio and carbon conversion increased
with increasing reactor temperature. These relationships are shown in
Figure 2 through Figure 5.
In conventional gasification processes, the phenomena occurring in
any gasifier are oxidation, reduction, pyrolysis and drying. The reac
tion zones for an updraft gasifier are illustrated in Figure 6. Several
reactions of importance in gasification are listed in Table 1. Starting
from the distribution plate of the reactor upward in the direction of
the gas flow, we find the following (6, 7, 8, 9):
i) In the "oxidation zone," the conversion with the
free oxygen takes place with the rate of reaction being high.
The oxygen pressure is high enough to favor COp formation.
The thickness of this zone may vary in magnitude from one to
ten centimeters.
ii) In the "reduction zone," several heterogeneous reac
tions occur simultaneously. These reactions are both endo-
thermic and exothermic. The high temperatures favor, kineti-
cally and thermodynamically, the C + CO2 = 2C0 and the
C + H2O = CO + Hp reactions which are highly endothermic.
Batchelder and Armstrong (6) in their coal gasification re
search indicated that steam diffuses simultaneously with
oxygen to the carbon surface, where it is decomposed by the
o -P cd (X
<u r-i O
o
Average Reactor Temperature, C
Figure 2. Effect of Reactor Temperature on H2/CO Mole Ratio (5)
\
6 .0
^ 5 .0 cu
M-
CJ 0 0
^ 3.0
to re
CD
/
/ CO
/
A/ / H,
/
/
/ .
2.0
V - — ' " ^
1.0
o r^ C.H Z'%
0 . 1 0 .2 0.3 Q.k 0.5
Oxygen/daf feed, l b / l b
Figure 3. Effect of Oxygen on Gas Composition (5)
8
T3 0)
Cd
rH 0)
03 Cd
Cd +» o
Average Reactor Temperature, C
Figure 4. Total Product Gas Yield from Sawdust (5)
\u
^
C o •H (Q U
^ O
o c o u cd o
1 0 0 -
Average Reactor Temperature, C
Figure 5. Carbon Conversion as a Function of Reactor Temperature (5)
10
Feed Product Gas
\L
Drying Zone (moisture driven off)
Endothermic
Degasification Zone (gas, tar and oil driven off)
Thermal Neutral
Reduction Zone (little or no oxygen)
Endothermic
Oxidation Zone (oxygen rich gases)
Exothermic
H2O {^) -> H2O (v)
Solid ^ Tar + Gas + Char
C + H2O - CO + H2
CO + H2O ^ H2 + CO2
C + 2H2 - CH4
C + CO2 - 2C0
H2 + 1/2 O2 ^ H2O
CO + 1/2 O2 - CO2
C + O2 - CO2
C + 1/2 O2 ^ CO
ash + unreacted carbon steam + air or oxygen
Figure 6. Major Processes Occurring in an Updraft Gasifier
11
Table 1. Important Reactions in Gasification
^^Rxn' ' J/9-'"0l
Reaction 298°K 1000°K
(a) CO + H2O = CO2 + H2 -41.2 -34.77
(b) C + 2H2 = CH4 -74.93 -89.95
(c) C + H2O = CO + H2 131.4 136.0
(d) C + CO2 = 2C0 172.6 170.7
(e) C + O2 = CO2 -393.8 -394.9
12
reaction: C + H2O ->- CO + H2. At the high flame tempera
ture, carbon monoxide is the main product of reaction and
little carbon dioxide is formed by the bimolecular steam
reaction. Product hydrogen and carbon monoxide diffuse into
the gas phase, where they are oxidized to steam and carbon
dioxide, respectively. Equilibrium is probably maintained
continuously in the reaction: H2 + 1/2 O2 = H2O. Hence,
during this period, the net effect of the steam reaction
is to augment the oxygen-carbon reaction. Steam present is
equivalent to an increase in the partial pressure of oxygen
at the high temperature at which diffusion controls both
reactions. Water-gas shift equilibrium is dynamically
maintained at the surface between the components steam,
carbon dioxide, carbon monoxide and hydrogen. Depending
upon the temperature and particle size, either diffusion
and/or chemical surface reaction, may be significant in
controlling the reactions.
iii) In the countercurrent operation of an updraft
gasifier a distinct "degasification zone" may be observed.
During the degasification, gaseous and condensable compo
nents (degasification gas, tar, decomposition water) are
evolved from the solid.
iv) The "drying" of the feed is affected by the hot
gases flowing upward. The temperature falls below 900°F
(482°C) and the reduction and shift reactions are "frozen".
13
A review of the investigations in the kinetics analysis are shown
below. Antal (10) used dry Whatman No. 1 filter paper as the feed and
presented this method according to Table 2 for his research.
e = gas residence time = ^ " ^f^ ^/Ps + V^v
where L was the total length of the gas phase section of the reactor,
£, which is l i s ted on Table 2, is a function of T for a steam flow of
0.34 g/min, T is the constant wall-temperature, a was the apparent
cross section area of the reactor; p , p , m , m were the densities
and mass flow rates of the steam and volat i le matter, respectively.
Figure 7 displays the dependence of gas production on gas-phase res i
dence time for various gas-phase reactor temperatures. Higher tempera
ture and longer gas residence time w i l l produce more gases. These re
sults indicated that biomass gasif ier should be designed to provide
for high heating rates and short gas residence time (5 seconds or less)
with gas-phase temperature exceeding 650°C. The gas-phase steam crack
ing reactions dominated the chemistry of biomass gasif icat ion.
Landeen (11) estimated the gas residence time for the SGFM reac
tor by plot t ing the reciprocal of the total gas flow rate versus the
reactor volume at eight di f ferent levels within the gasi f ier . For the
operating conditions: manure feed rate = 37.58 kg/hr, a i r feed rate =
7.05 kg/hr and steam feed rate = 16.51 kg/hr a l l at 577°C; the gas
residence time from the reactor bottom to the top was about 7.6 sec.
This also indicated that at higher temperatures, the gas residence time
would be somewhat lower because of the higher rate of gas generation.
14
Table 2. Parameters Used to Calculate Gas-Phase Residence Time (10)
T , °C w'
750
700
675
650
625
600
575
550
500
L = 29. ,2 1
a = 2.474
cm
cm
l^ cm
13.4
13.9
14.2
14.5
14.8
15.2
15.6
16.1
17.2
Effective Reactor Volume
3 cm
46.4
48.1
49.2
50.2
51.3
52.7
54.0
55.8
59.6
Effective Insert Volume
3 cm
30.2
31.3
32.0
32.6
33.3
34.2
35.1
36.2
38.7
Bulk volume of insert = 70 cm'
Length of insert = 31.1 cm
15
CO O
O
CO (T3
c o
•n-co s .
> c o CJ
c o
o
0.12
0.10
0.08
CO,
X o
— " A
/ ]/ _
0.06 i
0 0 8 10
0 8 10
0.5
0.4
0.3
0.2
0.1
0
CO
X
X o . #
o
_3_
_ A A
a I L
X J L X 0 1 2 3 4 5 6 7 8 9
Gas Residence Time, sec.
±
12
0.014
0.012
0.010
0.008
0.006
0.004
0.002
0
L ^2
X X •
a 1
A
1
a
x ^
X
o o
A
1 1 i 1 J i . . -
A
• •
Q
1 1
12
10 11 12
Figure 7. Gas Production vs. Gas Residence Time for Various Gas Phase Temperature: (a)500°, (A) 600°, (O) 650°, (•) 700°, (X) 750°C (10)
16
0 .08^ CH.
0.07
0.06
0.05
0.04
V) o
cu u
CO to ex CD
u.uo
0.02
0.01
0
. o
^ - - — A d
a
0 8 10
CO
(L) > o o
ctio
nal
<T3
u.
0.014
0.012
0.010
0.008
0.006
0.004
0.002
»
€
X
•X
- O
-
a
2 6
r
0 1 3 4 5 6 7 8 9
Gas Residence Time, sec.
12
10 11 12
Figure 7. (continued)
17
In measuring the rate of the steam-carbon reaction by a falling-
particle method, Dotson (12) calculated a solid residence time from the
relation t = w/F in which w = weight of carbon in the reactor at any
given time and F'= feed rate. The solid residence time was also esti
mated from Stoke's law and the linear velocity of steam in the reactor.
In these calculations, it was assumed that the particles were spheres
with a diameter equal to the average screen opening and that the steam
velocity along the axis of the reaction tube was twice that of the aver
age velocity (laminar flow). The solid residence time calculated in
this manner at 2000°F was 37% longer than the measured time. This dif
ference could be accounted for by deviation from the assumption regard
ing particle shape and size. This suggests that the particle shape fac
tor is yery important.
Yagi and Kunii (13) developed a method for estimating the average
residence time of solids in a fluidized bed. Figure 8 is a model of
such a fluidized bed. Taking the feed rate of particles to be FO, the
rate of overflow particles as Fl and that of carryover particles en
trained by the gas stream as F2, the total mass balance gives:
Fl + F2 = 3 FO (2)
where 3 = wt. fraction of unreacted feed. In Figure 8, KO(Dp), Kb(Dp)
Kl(Dp) and K2(Dp) are the size frequency distribution functions for the
feed particles, for the particles within the fluidized bed, for the
overflow particles and for the carryover particles, respectively.
By setting the weight of feed particles of diameter Dp equal to w,
the feed and discharge rates of particles with diameter between Dp and
18
Feed FO KO(Dp) Carryover
F2 K2(Dp)
Z>
Fluidized Bed w
Kb(np)
Overflow Fl
Kl (Dp)
Figure 8. A Model of Fluidized Bed, Particles with Wide Size Distribution (13)
19
Dp + d(Dp) can be expressed as the following:
feed rate of particles (number of particles per unit time)
NO = FO(Dp) d(Dp)/w (2)
overflow rate
Nl = Fl Kl(Dp) d(Dp)/(6w) (3)
carryover rate
N2 = F2 K2(Dp) d(Dp)/(3w) = W k Kb(Kp)d(Dp)/(3w) (4)
where k is the elutriation velocity constant
W is the weight of holdup in the bed
The mass balance for particles with diameter between Dp and Dp + d(Dp)
is shown in equation (5).
NO = Nl + N2 (5)
By substituting equation (2), (3), (4) into equation (5), an expression
for k is obtained.
_ B FO KO(Dp) Fl KUDp) ,g ^ W Kb(Dp) W Kb(Dp) ^ ^
The definition of the exit age distribution function for the over
flow particles is shown below.
El (Dp.ep) = (n/w)|lg£} exp[-{(n/w)flg} + wep] (7)
The average age, e"(Dp), of these particular overflow particle of
size between Dp and Dp + d(Dp) is shown in equation (8).
20
^0
>ep =00
0p NO de Ei(Dp,ep)de
I NO de Ei(Dp,ep)de. p
Then, the average solid residence time is obtained by integration and is
shown in equation (9). F(Dp) = average solid residence time
^ (9) (^VW)^^k
Substituting equation (6) into (9) results in an equation to calculate
average residence time of the particles.
e(op) = 3^ag}=e4o» no)
In conclusion, the average age (residence time in bed) of a wide
spectrum of particle sizes was found in all cases to be given by equa
tion (10) based on the following assumptions:
i) The fluidized bed operates at steady state with
continuous solid feed,
ii) Particle size remains constant during their stay
in the bed, but the average weight of the particles
becomes 3 times the fresh particle weight in com
pliance with the physical or chemical changes
(drying, pyrolysis, etc.).
iii) Complete mixing for particles with the same diameter
(the average residence time of a particle is usually
much longer than the time necessary for the mixing of
a particle after feeding).
CHAPTER 3
EQUIPMENT AND PROCEDURE
Experimental Apparatus
A schematic diagram of the SGFM pilot plant system used in this
study is shown in Figure 9. It consists of four sections: feed sec
tion, reactor section, tar collection section and downstream section.
Feed Section
Figure 10 shows the dimensions of the feed section which is com
posed of two parts, a hopper and a screw feeder. The hopper, which has
a capacity of approximately 40 lbs, is designed to prevent feedstock
bridging. The 1/4 Hp screw feeder, which has a variable speed control,
is manufactured by the K-Tron Co. (Model number S200). It was found
that infrequent bridging and leakage from the edge occurred within the
hopper making the feed operation yery difficult.
Reactor Section
Figures 11 and 12 show the schematic drawing of the reactor system
and the dimensions of the reactor. The reactor is constructed of sche
dule 40, 316 stainless steel pipe with 6" ID in the lower 5 ft and 8"
ID in the upper 3 ft. or disengaging section. The larger diameter of
the upper part is to minimize solids entrainment out of the reactor by
decreasing the fluid velocity and increasing solid-gas phase separation
The inlet gas distribution or plate of the reactor consists of a 3/8"
plate with 1/16" holes on concentric circles evenly distributed around
a 1" center port for char removal.
21
22
o
C3
CTt
(U
s -CD
PO
WE
R'
SU
PP
LY
<
2-3A T
t«
Js-i5-3/a •X
31 t t
\ /
10 l>
n. _ \ ^
^ ^ % ^ ^ ^ ^ ^
23
• ^
K-TRON
l A HP
U
]^ 1 5 - 1 / 2 " — ^
Figure 10. Dimensions of Feed Hopper and Screw Feeder
, f»-,«-.«
24
Screw Feeder fv=^=:
S igh t Glass
Bypass Gas
Power Supply
Figure 11. Reactor Section
25
l ' - 6 "
5' -1/2
6"
6"
6"
6"
6"
IL 5-1/2* z
900# Rf WN r i n g , s t a i n i a s s s t e e i
schedule kO p ipe s t a i n l e s s s t e e l
STD. WT. weld cap s t a i n l e s s s t e e l
schedule k-Q p ipe s t a i n l e s s s t s e l
concen t r ic r e a u c e r s t a i n l e s s s t e e l
schedule +0 p ipe s t a i n l e s s s t e e l
^ •12-l/'^-" - c o l l a r s t a i n l e s s s t e e l
900# ra i sed faced ••veld neck f l n g . w/"blind s t a i n l e s s s t e e l
Figure 12. Reactor Dimensions
26
An air-operated ram is used to prevent any bridging of the char in
the discharge line. A two-way solenoid and timer is used to control the
vertical movement of the ram and-allows: varying the cycle time.
Char is discharged into a collection vessel and removed manually.
The lower part of the reactor is heated with eight, 24" long elec
trical heaters, which were manufactured by the Lindberg Company (Model
50752, Type 77-1 CSD, temperature limited to 1200°C). These heaters
supply start-up and steady-state demands. Each of the heaters is rated
for 10 amps, 230 volts. The reactor is insulated with 5" of Cera-
blanket, which is a ceramic fiber manufactured by Johns-Manville, in
the inside layer and 1" of fiberglass which is manufactured by Owens
Corning, on the outside layer.
A U-shaped resistance preheater made of stainless steel tubing is
used to heat the inlet steam-oxygen mixture to 300-400°C. The preheater
consists of 16 ft of 3/4" tubing and 17 ft of 5/8" tubing. A Hobart
Model T-500-5422, 500 amp welding generator supplies power to the pre
heater by means of 3/4" braided copper welding leads. To prevent a
melt-down of the preheater piping, the contactor is controlled by a
Thermo Electric Model 32142-02-005 Mini Monitor Analog Latching High
Temperature Alarm triggered by a thermocouple clamped to the top of the
preheater.
The K-type thermocouples and sample ports are spaced at 6" inter
vals in the bottom of the reactor. Two additional ports in the upper
section, 1 ft apart, and one port at the exit are used to measure the
centerline temperature profile within the reactor. Reactor temperature
is controlled from a wall temperature measurement using a temperature
27
controller (Model 400, Thermo Electric Company). Temperature measure
ments of the preheater, reactor inlet, reactor exit, cyclone wall,
cyclone exit and impingers are also taken. All temperature measurements
are recorded by a 24-point recorder (Model 547, Leeds and Northrup
Company).
The cyclone is 4 ft tall with a diameter of 13" and is operated at
300°C to prevent condensation of tarry liquid products. Ten 1" by 6 ft
heating tapes (manufactured by Brisco Manufacturing Company) are wrapped
around the cyclone to supply heat. Each tape is rated for 115 volts,
384 watts with a maximum temperature of 482°C. Solids entrained by the
gas stream are by the cyclone removed manually.
Tar Collection Section
A diagram of the tar collection section is shown in Figure 13. The
gases leaving the cyclone pass through a two-stage impinger which is
operated at 110 ~ 140°C. These temperatures condense the tar but main
tain the water in a vapor state. Then, the gas proceeds to the double-
pipe heat exchanger where water is condensed and collected in the down
stream section. The major problem with the tar collection is the carry
over of extremely fine particles from the cyclone by the high rates of
product gas. This results in a viscous mixture that lodges in or be
tween the impingers and fouls the inside wall of the heat exchanger.
Steam can be injected through the impingers and the heat exchanger
after each run to remove any plugs that are formed. This procedure
satisfies the cleaning of the heat exchanger, but cleaning the imping- "
ers requires that they be disassembled and cleaned manually.
From Cyclone
?J r ~ i
V /- '
J
K.
JV H,0 ou t
u T
©
KJ
n e a t Zxchar.ger HjQ i n
To Downstream S e c t i o n
S teaa
7
^
28
Figure 13. Tar Collection Section
29
Downstream Section
The downstream section is shown in Figure 14. Gases and condensed
water from the heat exchanger first pass into a knock-out pot where most
of the gas and liquid are separated. The separated aqueous phase is
discharged through a float valve. The gases travel through a system of
three condensers where any remaining vapor is condensed. The last con
denser is packed with glass wool to remove any remaining fine particles
in the water-saturated gases before gas samples are collected and gas
flow rate is measured. Because the problem of plugging with the tar
aerosol on the gas flow meter (manufactured by Rockwell International),
a new method of determining gas flow rate has replaced the gas flow
meter. The new method consists of injecting a known amount of argon
into the product gas to calculate the product gas flow rate.
Operational Procedure
A run is initiated by heating the reactor to a desired, nominal
temperature while air is passed through the reactor. This heat-up time
is usually four to six hours. At the same time, the preheater, impingers
and cyclone are heated to the desired operating temperature. The steam
flow rate is set by condensing it in cold water for a set period of
time and measuring the increased weight due to the condensed steam.
Once the temperature of reaction is 100°C less than that desired, steam
is added to the air line through the preheater into the reactor. When a
thermal steady state has been reached (constant temperature to within
j 10°C), air is replaced by oxygen. Helium (the carrier gas in the gas
chromatograph) is purged to the feed hopper to prevent steam from
"1 r>-?--
30
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E 05 (U S> +J CO c 3: o o
CD
31
condensing in the inlet feedstock line. Feed is then started at a pre
determined rate. The feed rate is determined by collecting it in a
plastic bag for one hour and weighing.
One hour is usually needed for the reactor temperature to reach a
new steady state. Once the cyclone, impingers and char line are purged
completely, the material balance period begins.
During the material balance period (about 30-90 minutes), gas
samples are taken for analysis. Aqueous waste, char, cyclone fines and
tar are collected and their flow rates determined. Hold-up (the char
inside the reactor) is removed by the same line used for the char dis
charge.
The step-by-step start-up and shut-down procedures are shown in
Appendices A and B.
Analytical Procedure
Gas samples are collected in 250 ml gas collection bottles and
analyzed using a Carle Analytical Gas Chromatograph (Model ACG-lllH)
connected to a Spectrophysics Minigrator. Helium is the carrier gas for
all gas components except hydrogen, which has a similar thermal conduc
tivity to helium. Measuring hydrogen in a helium matrix will give low
sensitivity, non-linearity and peak reversals. Therefore, a hydrogen
Transfer System consisting of a silver-palladium diffusion tube and a
dual thermal conductivity detector are used. Hydrogen is transfered by
the Hydrogen Transfer System into a nitrogen carrier before measurement.
The carbon, hydrogen, nitrogen and oxygen content of the feed,
char and cyclone fines are determined by a Perkin-Elmer Elemental
32
Analyzer (Model 240B). The sample for elemental analysis is 1 to 3 mg
determined with a Cahn Electrobalance Model G. The moisture in the
samples should be driven off before the elemental analysis.
Moisture content of the feedstock, char and cyclone fines is de
termined by heating samples in a weighed crucible at 105°C overnight,
dessicating and then measuring the weight loss.
Ash content is also determined by heating the samples in an oven
at 950°C overnight, dessicating and then reweighing.
Particle size distribution of the feed, char and cyclone fines is
obtained by using the U.S. Standard Sieve Series (manufactured by the
W. S. Tyler Company) from No. 10 to No. 45 mesh.
Feedstock
Oak sawdust obtained from Missouri was used in this study. The
as-received feedstock contains 117% moisture on a dry basis. Then, it
is air-dried to approximately 17.6% moisture on a dry basis and screen
ed to -1/8 in. before each run. The elemental analysis performed on
the dried sawdust indicates that the wood contains approximately 50% C,
6% H and 44% 0 on a dry basis.
CHAPTER 4
DISCUSSION OF RESULTS
This chapter presents the experimental results and a discussion of
these results. The feed (oak sawdust from Missouri) was gasified with
steam and oxygen in the SGFM reactor operated as a countercurrent fluid
ized bed. Table 3 shows the operating conditions. Solids residence
time was varied by changing the ram rate at a constant reactor tempera
ture, oxygen-to-feed ratio and steam rate, except run 131 which was op
erated with no ram. In this study, the material balances and elemental -
balances are presented. The correlations between mean solid residence
time and product gas yield, H2/CO mole ratio and carbon conversion are
also reported.
Material Balances
During the material balance period, the input oxygen and helium
flow rates were measured by two separate Brook Type (1110-06 FIAIA)
rotameters. The rotameters were calibrated by a wet test meter. The
steam flow rate was controlled by two valves and set before each run by
condensing the steam in a bucket of cold water and measuring the increased
weight due to the condensed steam. Once the desired rate was reached,
the steam was injected into the reactor. A water rate was determined by
collecting all the condensed steam at a fixed period in the downstream
section. The input steam rate was considered accurate if the steam rate
and water rate were equal.
33
34
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35
In Yu's experiments (5), he indicated that the inajor source of error
was in measuring the gas flow rate. Yu suggested the use of argon as a
tracer to determine product gas rate, but the results were not satis
factory due to : 1) The gas rate was not constant during the material
balance period. 2) The argon entering the sample collection line was
not constant, because the pressure varied during the collection of gas
samples. 3) The argon rate determined by the wet test meter seemed lower
than what it should be. Therefore, for this study, argon is still to
be the tracer, but gas rate is adjusted by assuming 100% material bal
ance closure and calculating back to adjust the product gas rate. This
was checked by the elemental balances.
The elemental analyses shown in Table 4 were used to calculate the
elemental balances. Using a typical aqueous waste anslysis from Kao's
research (14), shown in Table 5, the carbon, hydrogen and oxygen content
in the aqueous waste was calculated to be 1.6%, 11.0% and 37.4%,
respectively.
After examining the elemental balances for run 126, 128 and 131,
it was noted that the aqueous waste rate was lower than expected. The
gas rate calculated from the material balance closure of 100% seemed
to be high because it gave a carbon balance of more than 130%. There
fore, the product gas rates of run 126, 128 and 131 were readjusted by
carbon balances. The readjusted gas rate was then used to calculate the
aqueous waste rate to make 100% closure of the overall material balance.
The results of the elemental balances indicate this method to be accept
able, except for the hydrogen balances. The results of these are shown
in Tables 6 to 9.
Table 4. Elemental Analysis of Solids and Gas
36
Run Number
Sawdust
% moisture
% ash
% C
% H
% N
% 0*
Char
% moisture
% ash
% C
% H
% N
% 0*
Cyclone fines
% moisture
% ash
% C
% H
% N
% 0*
Product Gas
% C
% H
% 0
Material from Reactor
123 124 125 126 128 131
17.38
0.86
49.80
5.85
0.18
44.17
8.24
7.83
62.09
4.26
0.20
33.45
20.22
15.77
72.61
2.34
0.57
24.48
40.01
4.84
55.T3
14.91
1.13
49.84
6.12
0.26
43.78
19.34
4.91
71.72
2.75
0.39
25.14
44.33
6.86
59.99
4.10
0.33
35.58
40.06
5.06
54.87
14.91
1.13
49.84
6.12
0.26
43.78
20.70
7.80
89.64
1.48
0.38
8.50
15.50
12.23
65.88
1.69
0.47
31.96
41.08
5.40
53.52
14.91
1.13
49.84
6.12
0.26
43.78
32.19
6.19
75.34
1.71
0.42
22.53
50.36
12.21
64.56
2.20
0.29
32.95
39.27
5.21
55.52
19.80
1.27
44.65
5.31
0.10
49.94
12.53
5.85
59.64
3.83
0.14
36.39
60.95
14.50
56.00
1.54
0.50
41.96
37.64
4.29
58.07
18.93
0.57
49.44
5.84
0.10
44.62
« • ^ ^
—
—
—
—
—
40.38
17.07
66.97
1.70
0.24
31.09
40.09
5.85
54.06
% ash 17.21 8.37 10.62 8.26 1 7 . 6 4
*% 0 = 100 - % C - % H - % N
37
Table 5. Waste Water Analysis from Kao (14)
Concentration (grams/1000 ml)
Methanol 3.382
Ethanol 0.394
Formic Acid 0.488
Acetone 0.660
Acetic Acid 26.150
Propionic Acid 6.161
Water 962.765
38
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It is thought the error occurring in elemental hydrogen balances
is due to: 1) The gas rate adjusted from material balance closure may
not be exact. 2) The feed rate varies and is only accurate to within
5-10%. 3) The samples taken for elemental analysis may not be repre
sentative of the whole material. 4) The amount of tar is neglected.
5) The amount of char, tar, cyclone fines and aqueous waste collected
during the period of material balance may not be exact.
Temperature Profiles
Figure 15 shows the temperature profiles within the SGFM reactor
for the six runs performed during this study. The upper section (great
er than 4.5 ft above the distributor) shows the effect of cold feed and
helium contacting and cooling the hot gas. The lower section (< 1.0 ft
above the distributor) shows the effect of steam and oxygen entering
the bed at lower temperatures. The temperature within the middle reac
tor is almost constant which is typical of a well-mixed, fluidized bed.
The average reactor temperature is the arithmetic average of the indi
vidual temperature measurements. This is consistent with previous
work.
Particle Sizes Distribution
Before considering the solid residence time of the different parti
cle size fractions in the reactor, it must first be determined how to
describe usefully the size distribution of solid particles.
Promesh paper (probability vs. mesh size) reported by Falivene (15)
and the method of least squares by Gauss are used to determine the mean
particle sizes for the inlet (feed) and outlet particles (char + cyclone
43
900
800
CJ o
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700
600
500
400
300 -
L
Lower Heater
V A
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t i
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a Run 124
A Run 125
V Run 126
O Run 128
X Run 131
Upper Heater
I I
Expanded Section
A
I
5 7 9
Thermocouple Number
i
11
fl X
Figure 15. Temperature Profiles of Runs 123, 124, 125, 126, 128, 131
44
fines). Figures 16 to 18 show the particle size distributions and Table
10 shows the mean particle size and particle size uniformity for each
run.
Because the feed samples taken for analysis may not be representa
tive of the whole material and because of the operational attrition in
side the reactor, the calculation of the mean particle size by the method
of least squares may introduce some error. Even though, the method of
least squares has a significant positive or negative deviation in the
measurement of mean particle size with uniformity, it is felt that the
assumption of constant particle size in the reactor should be acceptable
for calculating the solid residence time.
Determination of Mean Solid Residence Time
When the SGFM reactor was operated at steady state with sawdust
being continuously fed, it was a model similar to that of Yagi and
Kunii's gasifier (13). Therefore, equation (10) is used to calculate
the solid residence time for the wide spectrum of particles in the reac
tor.
First, it is necessary to know the solids hold-up (W) in the reac
tor for each run. Because of the temperature, the bed cannot be purged
immediately after the material balance period. The weight received from
the reactor is not accurate, because the char continues to react and
lose weight. The ash in the char can be used as the basis to evaluate
the actual solids hold-up as follows:
45
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47
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(sauas aAats puepue:is 'S'fl) Mssw
48
Table 10. Mean Particle Size and Particle Size Uniformity for Feed and Char + Cyclone Fines on Runs 123, 124, 125, 126, 128
Run 123
Run 123
Runs 124-126
Run 124
Run 125
Run 126
Run 128
Run 128
*Particle-size uniformity
Feed
Char
Feed
Char
Char
Char
Feed
Char
+
+
+
+
+
Mean
Cyclone Fines
Cyclone Fines
Cyclone Fines
Cyclone Fines
Cyclone Fines
Ma r— (from Promesh 'M
Parti c" (ym)
927
435
944
834
502
752
863
961
paper.
le
Fi
Size
gures
Particle-Size Uniformity*
16-18)
1.54
2.92
1.46
1.71
2.12
1.48
1.55
1.51
49
solids hold-up (W) = weight of char from reactor*
ash percentage in char collected at end of run /-j-jx ash percentage in char collected at steady state
The results are shown in Table 11.
Tables 12 to 16 show the residence time for each particle size
fraction and the mean solid residence time (MSRT). The calculation of
size distribution KO and Kl are referred to by Kunii and Levenspiel (16)
which is shown in Appendix C. The residence time for each particle size
fraction is not as good as expected because the larger particles should
have a longer residence time. The attrition between the particles as
well as the high velocity of the gas cause the deviation to make longer
residence time in the fine particles. However, if the actual solids
hold-up, based on the ash balance, is divided by the outlet flow rate
(W/3*F0) which is shown in Table 11, to calculate the mean solid resi
dence time, the same results are obtained. This indicates that either
method to calculate the solid residence time is acceptable.
Carbon Conversion
Carbon conversion is calculated by subtracting the unconverted car
bon in the solids leaving the reactor (char + cyclone fines + aqueous
waste) from the carbon in the feed and dividing by the carbon in the
feed. Most of the carbon consumed is converted to product gas. Fig
ure 19 shows the effect of MSRT on carbon conversion for this study.
As expected, the length of stay of particles in the reactor determines
the extent of reaction between the solids and gas. Thus higher MSRT
should result in more carbon conversion to product gas due to steam gas
ification of the char. For runs 123 and 125, the carbon conversion
50
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u. 1 -2
U. 13
^ D '33 \
T H
+ J
"T'U-Z
0 2 Z i i * * ^
N -^ 1*0 CO O "!• CO llT'iJ-'J-.-^ LL-^^O
O
•70 S3
* ZU"!" cr'33 cr < r > - » ^ •?•. c-4 K h - .
O UJ .'?-4 03 1:0
'?-. I! I! I! 1! v LU Ci O'^C-4 \ C C
: ^ 3 } e : { e * : 4 e * # * « * « * * * > } t * : ) e * j ) c * : 4 c * * * j } e * « j S t s ^ : } : : ^ e s } e i | ' i i " | j f | | ^ j j || LLJ O
o o o o o o ^ rs o ij-i ,-, 5; ,-.._. o o CO '^ 'j"- o . cr o c-- ij" c-4 O S3 «*• ' ^ O CO N |j"J '70
C-4 T H
C3 T H
'70 Ci
CO •N
T H
'70 •> T H
<» CO
c-4 C-4 T H T H TH T H
O O c-4 ^ s3 '70 O • ^ 1 ^ T H T H T H T H I"-J C.V
in c-4
UTi ' . ' I
l i . l
cr
p LLV;LU U-O-V^f: Z
- J - ^ 2 i—• _ UJ C4II ! -
C I C C . J Q L U ' - ' LLI<II«I_J + -:I . UJX>- 'Z'THWC'*' LL'-"I:IXLLZUJ
W H H C C -J -J_J-J l ! t -
ccccc -z •_"Z' CI 'Z 'UJUJ LU
56
^
to &. cu > o
C_) c o
. Q s -03
C_>
100 -
90
80
70
60
50
40
30
O ^ " o.
20l
720-740°C
.23 lb 02/lb Feed
I 1 1 0.0 2.0 4.0 6.0 8.0 10.0 12.0
r^an Solid Residence Time, min.
-ff-
O 53 min.
Figure 19. Effect of MSRT on Carbon Conversion
57
appears to be low. The result from run 131, which is operated with no
ram (longest MSRT), shows the highest carbon conversion. This indicates
the scatter in the data is probably due to experimental error.
Product Gas Yield and Compositions
The purpose of gasifying oak sawdust with steam and oxygen is to
produce gas suitable for methanol synthesis. This study is to investi
gate the effect of solid residence time on the product gas yield and
H2/CO mole ratio. A high Ho/CO mole ratio is desired because it needs
to be 2.1-2.2 for methanol synthesis. Therefore, besides the gas yield,
the gas compositions are also significant. As expected, the product gas
includes H2, CO, CO2, CH^, C2H^, C2Hg, C2H2, etc. for each run. Table
17 shows the product gas yield and mole fraction of each component. Fig
ure 20 shows the effect of MSRT on product gas yield. It indicates that
the product gas yield increases dramatically during the first 2 minutes
and continues to gradually increase when the MSRT is more than 2 minutes.
This is due to steam gasification of the char.
The effect of MSRT on gas compositions is shown in Figures 21 to 24.
In combination with run 131 data. Figure 21 shows that carbon monoxide
and hydrogen increase with MSRT and carbon dioxide remains constant after
a MSRT of about 3 minutes. Methane and ethane yields increase with in
creasing MSRT, but other hydrocarbon yields remain constant at varying
MSRT's.
Because this study was conducted at a high temperature C720-744°C),
which favors high gas production in a short reaction period, both carbon
dioxide and carbon monoxide are formed as primary products of the
58
131
128
.277
CD CM
.727
CO CM
.648
CM CM
.824
p^ CM
.970
CM
.278
CM
CM
CO
O
436
o
205
o
203
•
o
886
13.
455
CD
287
30.
964
32.
LO
•
o
171
•
o
CM rs. CO
• CM CM
CO
«;r CM
CO CO
ions
si
t mp
o
o CJ
T3 OJ
•o ^— cu >-
to OJ CD
126
125
124
123
.292
CO CM
.180
Ps CM
.341
to CM
.598
CM
.160
CO CM
CM LO CM
CO CM
.395
CO CM
.262
CO CM
CD CM O
CO
.219
CO
.490
CO
.064
CO
,685
o 634
•
o
.523
o
,591
o
1 t
145
•
o
,285
o
922
o
364
13.
486
CM 647
CM
120
CM
250
30.
082
30.
320
34.
895
LO CO
221
•
o
1 1
, 244
o
LO CD
rs.
ro CM
CO
ro CM
CM
r^ ro CM
0 0
CD
ro CM
0 0
CD
CT>
o
cu
03
CU p—
o E
XH
"s^ Xi
XJ cu cu
t i -
4-03
E CD
& . CU
s-cu
Xi E 3 Z
C 3
OC
•"" ^ —'' CM
^
--^ ^ ^—^ CM
O CJ
^« — «;r
I CM
CJ
^ ' — CO
CM CJ
5^
CM
CM C_)
^.^ &« >_ "^
3: o
«—s ^
o CJ
^_^ ^
+ CO CJ
•
^^ o • CD > <
T3 ^ ^ CU •r^ >-
(/) 03 CJ3
K2|.
1.1
59
1.0
I 0.9
ZO.8
5 0.7
01
/ O h /
/ L/ /
U
>-
^ 0 . 5
u - i 0.4 o a.
f
0.3
0.2
0.1
0.0
O :experimental data
:corre lat ion from SAS
3 4 5 5
Mean Solicj Residence Time, min.
Figure 20. Effect of MSRT on Product Gas Yield
60
0.5
- ^ - H2
*•*• CO2
• " A - CO
• o cu cu
03
CD
s-cu
XJ
cu • r ->-
CM
o o -a c ro
O CJ
CM
0.4 -
^ -
> -.'iT
— CO
0.3
y 2'
0.2
0.1
H-
o a
6
0.0 1 2.0 4.0 6.0 8.0 10.0 12.0
Mean Solid Residence Time, min.
-ff
Figure 21. Effect of MSRT on H2, CO, CO2 Yield
61
cu cu
03
E cr s-cu
T3
CU •r ->-
CM CJ
CJ
0.20
0.15
0.10
0.05
0.00
CH.
• ^ ^2^4
CH
/
> /^
/
OO o
C2H,
1 I. 1 0.0 2.0 4.0 6.0 8.0 10.0 12.0
Mean Solid Residence Time, min.
Hh
Figure 22. Effect of MSRT on CH^, C2H^ Yield
62
0.010 -
XJ cu cu
OJ T3
E CD
& . CU
-a r— CU
CM IC
CM O
CO
zc CM
CJ
0.008
0.006
0.004
0.002
0.000
^ hh CrtHrt
/
/
A . /
/
- o o
1
'zh
CpHp
J. 0.0 2.0 4.0 6.0 8.0 10.0 12.0
Mean Solid Residence Time, min.
-ft-
Figure 23. Effect of MSRT on C2Hg, C2H2 Yield
63
OJ OC
cu o
o o CM
1.0
0.9
0.8
0.7
0.6 -
0.5
0.4
^
/
0.3
0.2 L I ' I
0.0 2.0 4.0 6.0 8.0 10.0 12.0 •if-
Mean Solid Residence Time, min.
Figure 24. Effect of MSRT on H2/CO Mole Ratio
64
combustion and their ratio is a function of temperature. Due to equili
brium, carbon dioxide predominates at low temperature and carbon mono
xide predominates at high temperature. This relationship was partly
shown by Lambert's (17) work on the low temperature oxidation of coco
nut charcoal. He found, when using mixtures of oxygen with carbon mono
xide or nitrogen, that at 512°F the gas-phase oxidation of carbon mono
xide had a negligible rate. Carbon monoxide acted merely as a dilutent
and its concentration was slightly increased. The primary product was
carbon dioxide. However, as the temperature was raised to 836°F the
oxidation of carbon monoxide completely overshadowed the primary reac
tion at the carbon surface. Then finally in the combustion zone, the
temperature was so high that the carbon monoxide was the predominant
product.
2C + O2 - 2C0
2C0 + 0 2 " ^ 2CO2
The carbon-steam reaction (C + H2 -»- CO + Hg) favors production of
hydrogen and carbon monoxide. The components CO, CO2, and H2 are in
equilibrium according to the relative amounts of carbon, oxygen and
steam. Because this study was conducted at constant steam and oxygen
rates, as the MSRT is increased, the oxygen is depleted so fast that it
contributes very little to additional CO2 formation. The increasing
carbon conversion leads to a gradual increase in formation of CO and H2
The relation between product gas yield and MSRT was correlated by
SAS (Statistical Analysis System). A multi-regression analysis of the
65
data gives the following non-linear equation based on the assumption
that the gas yield is 0.22 i/g daf feed (oxygen rate) at MSRT = 0.0.
G = A * £n(D + R ** B) + C, 1.08 < R < 7.44
The percentage root mean square deviation (% RMS Dev.) is 4.50%.
n G.(correlated) - G.(experiment) o ,,« % RMS Dev. = [ E (-5 r l r. n x lOOVn]^/^
\j_1 G. (experiment) ' ' -•
where G = Product gas yield, liter/gm daf feed
R = Mean solid residence time, min.
A,B,C.D = Parameters which are a function of the operational conditions
(temperature, oxygen and steam feed rates)
For the temperature, oxygen and steam feed rates of this study, the
values of the parameters are shown below.
A = 0.469 C = 0.870
B = 0.223 D = 0.25
This form of equation was chosen because of the following:
1) I t f i ts the data very well as shown in Figure 20.
2) As expected, carbon conversion increases non-1inearly with
mean solid residence time.
3) Because rapid degasification occurs at a short MSRT, the pro
duct gas yield should have a dramatic increase and continue
to increase gradually.
4) I t is then reasonable to see that the calculated MSRT of
run 131 using this correlation is about 52.4 min.
66
Effect of Other Variables
Using the results of Yu's work (5), the effects of other variables
such as temperature, steam/daf feed ratio and oxygen/daf feed ratio on
the product gas yield, H2/CO mole ratio and carbon conversion were also
considered. Table 18 shows the data which were correlated by SAS. The
following correlations were obtained.
G = 0.30498 * Jln(0.25 + R ** 0.223) + 0.00295 * T +
0.67112 * 0 - 1.40366
Q = -0.01254 * R + 0.00731 * R ** 2.0 - 0.00050 * R ** 3.0 +
0.00390 * T + 4.32750 * 0 - 7.20199 * 0 ** 2.0 + 0.00244 * S +
0.55093 * S ** 2.0 - 2.81186
CC = 0.05227 * R - 0.00320 * R ** 2.0 - 0.00017 * R ** 3.0 +
0.00011 * T + 1.87376 * 0 - 3.20182 * 0 ** 2.0 - 0.13884 * S -
0.18269 * S ** 2.0 + 0.46111
The % RMS Dev. is 8.50%, 11.67% and 3.61%, respectively,
where G = Product gas yield, liter/gm daf feed
Q = H2/CO mole ratio
CC = Carbon conversion, %
R = Mean solid residence time, min.
T = Temperature, °C
0 = Oxygen to dry, ash-free feed ratio, lb/lb
S = Steam to dry, ash-free feed ratio, lb/lb
These above correlations of the product gas yield and H2/CO mole ratio
are similar with those proposed by Yu if R is constant. Table 19 shows
the errors in these correlations which indicates these correlations f i t
the data very well. Because the operating conditions for all these runs
67
'X) ^ C D C D O ^ C M f * ^ * '~ ^ S ; : 3 : ' ^ ^ C M C M »—• I— 0 0 0 0 P s r - ^ r > ^
C>0 <: CO
I— o O O r -
0 0 CM r—
o r—
o
r— r— 0 0
^ CO 0 0
CM CM p s
CO 0 0 CM
CD
«;r 'd-
LO 'sr
o o O O CM
LO 0 0 CD CD r—
c o
• r -CO s. cu > c o
CJ
c o
Xi &. OJ
C-J
• o
OJ
" ^ CO CM Ps. ^e
— r-> S ! ^ ^ ^ < J f— O f s . 0 0 ,— ^ f.^
I— I— o
CO LO O ^ I — O o CO r*v
• . . «— o o
O O r -
v o CM r^ ^t CM CO CM CM P^ r - ^ p^
. . . O O r -
C D O
' ^ O r— LO CO O 0 0 CM 0 0 O Ps ^ CO CO r—
CM -k ^ CM CO r*^
o o o O r—
ps. CM r—
0 0 CO 0 0
CO
^ CO
0 0 r^ p ^
"sT
<:P CO
00 <*•
o 0 0 CO
•K CM r^
o o o O O r -
CO CM r ^
o 0 0 f —
CD CO 0 0
r^ CD 0 0
^ ^ P s
CD CO CM
O 0 0 CO
CM CM
I— o o O O LO
LO CM
< —
o o r—
^ o CD
0 0 p s 0 0
o «sr p ^
CD CO CM
O CO CO
«d-•sT
I— o o o o
03 OC
Q)
O
O O
CM
zc
'aj > -
to fO
CD
C D
c •r— PsJ
03 C
o ^-
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a
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^ CM CM
L O O ps.
CM 0 0 CO
CO O p ^
CO CD CM
•K CM r>
CO
o
O O I—
CM CTt CD O «:d- P>s * CO < ^ t o r^ O CO CM CO CM CO CO r— ^ pv.
o o O O I—
^ CM
<—
o "xp CD
0 0 CO
r^
"sr CD r s
o CM
(^
CD CO CM
O 0 0 CO
0 0 O
o o O O I—
ro CM i~~
o cn o
LO 0 0 CO
0 0 to p-s
CM ro !>>.
rs. r>. CM
CO CO «:*•
CM p«s
o o o o .—
CM CM 1 —
ro ro < —
CO ps. ps.
«^ «!3-rs.
CO
^ r^
,— CM
^
^_ r^ ^
* CM rs.
r— O O O O I—
CO
CU
x> 03
"CJ CU CU
OJ
CD
S- O CU • ! -
4-i •»-> • r - (O _ J QZ
Ol Xi
3 -p-
c o
•r-to
cu > c o
CJ
CJ o
cu
03 "O S- CU cu cu
c 3
OC
- o I— r - O cu s:
sz >- o o
t j j a to *v^ s_ OJ CM 03
CU 3 Z CJ
CU
o •M a OJ cu
OC
Xi
Xi
-o cu cu
t » -
03
-o c cu C D > > X
o
03 • ! --O E
E OJ cu
•M OC CO
s-cu .o E 3
2r
c 3
OC
• o
cu cu
M-«+-OJ
T 3
E CD
« t
S-
cu 4-i •r— __J
9S
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CU • F -
>-CO 03
CD
o •r— • M OJ
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zc
sz o
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%. CU
> sz o
C J
c o
Xi
u 03
C J
C J o
w*
cu 3
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Xi
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ro CM
SZ 3
• p -
CU 4-> OJ L .
E OJ %-
cu E 03 CO
cu
cu 3
cu CO
to
cu
3
OJ
>
to •r-I—
68
CO ^ o CO O r— I— r— O
• • • p - p— CO
cy» r^ ^ CO T3-0 0 0 0 ^
. . . O O p -
cy» LO p - t o CO 0 0 r->. CO
. . . O O IS..
00 CM r^
O LO p— CO CO O O '^
r— O 1— r>* LO
00 r o
^ CO CO CM o 00 CO CO
p— r— CM O o to O O p -
LO 0 0 cy> CD O I— r - CM CD
• • • p - p - o
r o «?r ' d - 'ia- CO O CD LO
• • • p - O CD
CD p— f—00 CM
o CD rs. * • • r— O CO
LO O p— CO CO CO CO ^
«?r o^ ^ LO CM ps. rs, o . . .
O O CM
O <D o 'r o 1^ rs. o
127
126
00 CM CO CT> 00 00 CO CO • • •
O O ps
O CM 00 CO O
CO O ;r CO ^ CO CM 00 • • •
O O ':a-CM
CD LO CO r>s CO 00 00 ps.
00 t— P^ «^ ps. p>«- rs. ps. • • •
o o ^
p^ cy> O^ CT>, CT> 00 00 1—
O O CM O O P>. r— f— <d- O O O o o o
to sz o •p—
4-> 03
CU
s-s-o CJ
c • • "
to s-o i~ i-LU
CD
CU
J3 rO 1—
CD O r—
00
o r—
CO
o p ^
^ LO O 00 CM p«. rs lo • • •
O O r— r— 1
P-.. CO CO P^ CM r^ rs CO • • •
o o to 1
CM O CO «;i- p— CO PS. O • • •
o o ps. i
O CO 00 VO p-«!:r <* LO • • •
O O CO
"^ CM CM ps. CO CM CM CO • • •
O O l — CM 1
CD CM <;r CD «d-CM CM p— • • •
o o r 1
f— CO CM p- IS. CO CO 00 . . .
o o o
to CM O CM CO ps. rs. «;f • • •
O O CVJ
CD p— LO r ^ p>» CO CO 00 • • •
O O P -
LO CM
"^ CM r—
CO CM r^
CM CM
•
o o o 'r 00 p— 1— CO • • •
1— p- CO 1
O CM "sT P^ O . CD CD «^ • • •
O O CO 1
O 00 o> CO CO
o o rs • • •
p - I — ^
CO CO CO P>s CM
1.1
1.1
-3.8
«^ CM O O CO CD CD p— • • •
o o o
00 CO CO r— CM 1^ Ps CO • • •
O O CO
L O r s 00 O CD CO 00 P^ • • •
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i
00 CO rs CO 00 p>. rs. CX5 • • •
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00 00 r^ ps 00 00 • •
o o
'sT r^ CD 00 ps IS, • •
o o
00 CM LO <Tt r^ ps • •
o o
« o
«*
o •
o
CM 00 •
o
^ ^ •
«;r 1
'^ '^ 00 P^ ps, • •
o o
0.5
• ^
o o p —
X
nta
l im
e s-cu
'Exp
-o cu 4->
o •^ x> cu s-Q.
03
fO
s. <U "O
Xi I— E cu 3 -p-
z >-c to 3 03
OC CD
CU XJ CU 03
•p- O O S- -p- i -CU XJ i . Q . CU LU X s-
LU Qu ^ ^
••-> 03 03 -M
OC SZ-XD g j CU
CU E 4-> S-r— .p- O O O S- ' r - i -
s : CU XJ s-O . cu LU
O X s-
CJ LU C^ ^ ^
CM
c o •r— (O s-cu > o
CJ c o
Xi i-03
O
03 - P C XJ CU CU E •<-> S-
•p- o o S- -r - S.. cu X J s-Q . cu LU X s-
LU CI. S ^
03 +•> c -o gj CUOJ
S- E 4-> S-CU XJ - r - O O
XJ p - S_ -p- i . E CU CU XJ S-3 T - C2. OJ LU
Z > - X s-LU Q_ ^
c to 3 OJ
OC CD
OJ OC
03
X» CU CU
CU E 4-> I— •r- O o i - - i -s: cu "o
CL CU O X s-o LU a .
CM
s-o s .
o
CO i-cu > c o o c o
Xi
OJ
o
03 • M C X3 m CU E 4-> i -
•p- o o
cu "O s-C3. cu LU X J -
L u a . ^
s-cu CL X
LU
II
& -o s-u
OJ
69
were within the ranges of average reactor temperature = 644 - 816°C,
oxygen/daf feed = 0.104 - 0.421 lb/ lb, steam/daf feed = 0.298 - 0.471
lb/lb and mean solid residence time = 1.08 - 7.44 min. the correlations
are valid only over these ranges of conditions. Additional runs should
be performed to extend the range of application.
CHAPTER 5
CONCLUSIONS AND RECOMMENDATIONS
Conclusions
1. Mean solid residence time varies when the ram cycle is changed,
The larger the ratio of ram up to ram down, the longer the
MSRT.
2. Carbon conversion, product gas yield and H2/CO mole ratio in
crease with longer MSRT, but its effect is not wery prominent
after 2 minutes.
3. In order to increase the H2/CO mole ratio, gas yield and
carbon conversion at a given temperature, the reactor should
be operated with no ram, which would increase the MSRT.
4. Using the ash percentage in char as a basis for calculating
the real solids hold-up is reasonable.
Recommendations
1. More runs at different operating conditions should be perform
ed to confirm these results. A study should be performed to
determine whether the predominant factor is temperature or
MSRT.
2. Study the optimum conditions with no ram in an effort to get
a higher product gas yield and higher H2/CO mole ratio.
3. Resolve the problem in measuring the product gas flow rate.
An accurate reliable gas flow meter should be used. Fine
70
71
particles and aerosol in the gas stream should be removed be
fore entering the gas flow meter. A venturi scrubber follow
ing the cyclone will probably remove the fine particles and
tar aerosol.
LITERATURE CITED
1. J . E. Halligan, K. L. Herzog, H. W. Parker, "Synthesis Gas from Bovine Waste," Ind. Enq. Chem. Proc. Pes. Dev., 14(1), 64-69 (1975). ' —V / ,
2. H. G. Hipkin, D. J. Basuino, "Syngas from Manure - A Conceptual Plant Design," Final Report (1978).
3. S. R. Beck, M. J. Wang, J. A. Hightower, "Pyrolysis of Biomass in the SGFM Process," 72nd Annual AIChE Meeting, San Francisco, CA, Nov. 25-29 (1979).
4. S. R. Beck, M. J. Wang, "Wood Gasification in a Fluidized Bed," Ind. Enq. Chem. Proc. Des. Dev., 19.(2), 312-317 (1980).
5. J . F. Yu, "Process Variable Study on Steam-Oxygen Gasification of Wood," M.S. Thesis, Texas Tech University (1981).
6. H. R. Batchelder, R. M. Busche, W. P. Armstrong, "Kinetics of Coal Gasification," Ind. Eng. Chem., 45^(9), 1856-1878 (1953).
7. Vaclav Biba, J i r i Macak, Erhard Klose, and J i r i Malecha, "Mathematical Model for the Gasification of Coal Under Pressure," Ind. Eng. Chem. Proc. Des. Dev., 17(1) , 92-98 (1978).
8. Heeyoung Yoon, James Wei, and Morton M. Denn, "Feasible Operating Regions for Moving Bed Coal Gasification Reactors," Ind. Eng. Chem. Proc. Des. Dev., 18(2) , 306-312 (1978).
9. Solar Energy Research Inst i tute, SERI/TR-33-239, Volume I I , Chap. 6, July (1979).
10. M. J. Antal, J r . , "The Effects of Residence Time, Temperature, and Pressure on the Steam Gasification of Biomass," 177th ACS National Meeting, Honolulu, Hawaii, April 2, 1979; "Biomass as a Nonfossil Fuel Source," published by American Chemical Society (1981).
11. B. C. Landeene, "Internal Reactor Gas Profiles Using Bovine Residue as Feed," Unfinished M.S. Thesis, Texas Tech University (1977).
12. J. A. Dotson, W. A. Koehler, J. H. Holden, "Rate of the Steam-Carbon Reaction by a Falling-Particle Method," Ind. Enq. Chem. Enqr. and Proc. Dev., 49(1), 148-154 (1957).
13 S. Yagi, D. Kunii, "Fluidized Solids Reactors with Continuous Solids Feed," Chem. Eng. Sc i . , 16_, 364-391 (1961).
14. C. H. Kao, "Extraction of Waste Water from Biomass Gasification," M.S. Thesis, Texas Tech University (1980).
72
73
15. P. J. Falivene, "Graph Paper for Sieve Analyses," Chemical Engineering. Feb. 23 (1981).
16. D. Kunii, 0. Levenspiel, "Fluidization Engineering," published by John Wiley & Sons, Inc. (1969).
17. J. D. Lambert, "The Oxidation of Carbon," Trans. Faraday Soc, 32, 452, 1584 (1936).
APPENDIX A
SGFM START-UP PROCEDURE
1. Check that all valves and connections are properly set
- first vent by-pass should be open - steam, air, oxygen, helium and water should be off - gas line to gas meter should be closed - reactor expansion joint should be loose.
2. Start air to reactor and set rotameter to maximum. Close valves which are venting to operating area. Check for leaks.
3. Turn on power to panel and start temperature recorder. Turn on steam to impingers. Set the temperature alarm to 950°C.
4. Start lower reactor heaters:
Step 1 = Turn all circuit breaker switches off; Step 2 = Turn on main power and set controller to 800°C; Step 3 = Turn on pair of two circuit breakers on lower heaters.
5. Start upper reactor heaters:
Step 1 = Turn all circuit breaker switches off; Step 2 = Turn on main power and set controller to 800°C; Step 3 = Turn on pair of two heaters.
6. Switch on preheater:
Step 1 = Turn on preheater panel switch; Step 2 = Monitor and adjust the preheater temperature to
400-500°C.
7. When reactor temperatures are within approximately 50°C of set point values, tighten expansion joint.
8. Open second by-pass lines and close first by-pass line. Turn on water to gas heat exchanger.
9. Adjust water flow to the upper impingers to set the impinger temperature to 105-120*'C.
10. Inject steam to the reactor. Adjust the steam rate to the desired value by measuring the water"rate from the lower impingers.
11. Open lines to gas meter and check meter for operation. Close lines to gas meter after checking.
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12. When reactor temperatures are in steady state, switch air to oxygen and adjust oxygen feed to the desired set point.
13. Open f i rst by-pass and close second by-pass.
14. Set feed controller to the desired value. Turn on helium to feed hopper. When the hopper pressure is greater than the column pressure, open the valve between the reactor and hopper. Adjust helium flow and make sure the hopper pressure is always greater than the column pressure.
15. Turn on power to ram timer. Start solid flow and check sight glass for feed. Monitor and adjust the feed hopper pressure. Watch the lower reactor temperature closely.
16. Remove char and cyclone fines ewery 10 minutes.
17. Monitor and adjust reactor, preheater, and impinger temperatures. Monitor and adjust helium flow. Check sight glass for feed all the time.
18. Adjust the reactor temperatures to the desired values by changing the set point of the controllers and turning some of the circuit breaker switches on or off.
19. When the reactor temperature is in steady state, open second by-pass and close f i rst by-pass slowly. Monitor the reactor pressure when closing f i rs t by-pass.
20. After 15 minutes, start material balance for desired period
- remove all the water, char, cyclone fines, tars before starting material balance
- check sight glass for feed all the time - monitor and adjust the operating condition - measure the product gas rate - collect 2(3) product gas samples for GC analysis - collect waste water and chars eyery 4 minutes - collect cyclone fines every 10 minutes - at the end of material balance, collect all the waste water,
cyclone fines, chars and tars - mark on the temperature recorder the beginning and the end
of material balance.
APPENDIX B
SGFM SHUT-DOWN PROCEDURE
1. Turn off solid feed and oxygen flow.
2. Close the valve between reactor and feed hopper. Turn off helium flow and turn on air to reactor at maximum rate. Leave the drain valves of lower impingers open.
3. After 5 minutes, turn off reactor heaters. Switch to first by-pass Purge tar traps by venting through all drain valves.
4. Wait 15 minutes, turn off steam flow to the reactor and loosen expansion joint.
5. Wait another 15 minutes, turn off preheaters and ram.
6. Shut off steam, or water to upper impingers; leave drain valves open. Turn off water to heat exchanger.
7. After all reactor temperatures are less than 50°C, turn off air. Remove char and cyclone fines.
8. Normal maintenance and other items:
a. clean impingers and heat exchanger by steam b. check calibration on feeder c. clean feed sight glass d. check air leaks on ram e. dump cyclone and char hopper f. replace glass wool in filters g. refill feed hopper h check preheater connections every second run i. check the quantities of oxygen, helium and nitrogen left j! check all the connections, thermocouples, valves, and
other items k. clean-up, sweep-up, etc.
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APPENDIX C
EXAMPLE CALCULATION OF t€AN SOLID RESIDENCE TIME (MSRT), RUN 123
KO(J) = (SUMWIO) . ^^"^^"^J . SUMWTQ(J) - SUMWTOfJ-1) ^"^ ^ SIZE ' j {SIZE)j S i / h ( j ) ^ SIZE ( j -1 ) '
K1(J) = ( ^ ™ - ) = i !!J^l^-J - SUHWTUJ) -SUMWTKJ-n ^ SIZE ^j (SIZE)j SlZt ( j ) - SIZE(J-l)
from equation (10)
Average Solid Residence Time (J) = ASRT(J) = ^ * I I * ) , , 3 K 0 * K 0 ( J ) '
J = No. of screen, from < 10 to 45
from equation (11)
a
,a" Solids Hold-Up = W = 0.0265^ * R ^ ^ ^53 g ^ 26.37
7.83^
Konn - 1'QQO • o«979 _ ^ .^. \(Uiu - i-ooo - 0.959 ^ ^^^ •^^^'^^ - 2.200 - 2.000 - °-^°^ ' ^^^^^^ " 2.200 - 2.000 = ^'^^^
i\<zDT(ii\ - 26.37 * 0.155 - ? c-j M5K1UU 0.174 * 88.2985 * 0.105 " ' "
^.^(.r.y. _ 0.979 - 0.937 _ ^ ,.,, ,,w,f.x _ 0.969 - 0.922 _ ^ , . 7 ^^^^^^ ' 2.200 - 1.680 " °-^^^ ' ^^^'°^ - 2.000 - 0.1680 " ^'^^^
AcjRTnn^ - 26.37 * 0.147 ^ . g^ MbKiuu; - Q -,7^ * 88.2985 * 0.131 ''^'^
i/n Q^ - Q-937 - 0.869 _ ^ « „ ,,WQX _ 0.922 - 0.853 _ ^ „.^ •^0(9) - 1.680 - 1.410 - °-2^2 ' ^^^^' " 1.680 - 1.410 ' °-256
Ac:RTfq) = 26.37 * 0.256 ^ , j. i\:iKiK^) 2 . 1 7 4 * 8 8 . 2 9 8 5 * 0 . 2 5 2 ' • ' ^
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K0(8) = 0.869 - 0.785 _ ^ ^r.r, |,w«v . 0.853 - 0.801 . ^ ^^ ^^^^^ 1.410 - 1.190 " ^•^^^ ' ^^^^^ • 1.410 - 1.190 ' °-2^^
AC:DT^O\ _ 26.37 * 0.236 , . ^ '^^^ ' 0.174 * 88.2985 * 0.382 ~ ' '"^
uQ/yx _ 0.785 - 0.639 _ r. , . 0 ^w,v _ 0.801 - 0.723 ^ - , , • " ^ " 1.190 - 1.000 " °-^^^ ' ^^^^^ - 1.190 - 1.000 " Q- ^
ASRTr7) = 26.37 * 0.411 _ . MOKiu; 0.174 * 88.2985 * 0.768 " ^'^'^
Knrfil - Q»639 - 0.492 _ ^ p.. , . _ 0.723 - 0.660 _ ^ . Q . • ^ ^ " 1.000 - 0.841 - °-^25 , Kl(6) - ., QQQ _ Q g ^ - 0.396
ACDTf6 . 26.37 * 0.396 ^^ M5KKo; 0.174 * 88.2985 * 0.925 " ^''^
i/n/c\ 0.492 - 0.372 ^ QQ^ ,,.,/C-X _ 0.660 - 0.586 _ ^ ceo ^Q( > = 0.841 - 0.707 = Q^ ^ ' ^1(5) - 0.841 - 0 .707 " ^'^^^
26.37 * 2.582 ^^•^^(^^ ^ 0.174 * 88.2985 * 0.896 = ""- ^
^Cf)(A) - 0.372 - 0.251 ^ , 034 KU4) = 0-586 - 0.518 ^ Q ^g^ • " ^ " 0.707 - 0.590 '•"'^^ ' ' ^^ 0.707 - 0.590 ^•^^'
ASRT(4) = '1^74 * 88^2985 * 0.034 = ''''
^nf^\ 0.251 - 0.073 _ n 7t:7 K U 3 ) = MI^-^L-MIi = 0 847 • 0(2) " 0.590 - 0.355 " ^'^^^ ' '^^ 0.590 - 0.355 ^'^^^
,^^^f^^ 26.37 * 0.847 , -i go ASRT(3) = 0.174 * 88.2985 * 0.70/ ''^
.,^,^, 0.073 - 0.000 _ n onfi Kl f2) = 0.319 - 0.000 ^ « ggg '<0(2) = 0.355 - 0.000 ' ^'^^^ ' ' ' ^^ 0.355 - 0.000 "'^^^
„ „.r/o - 26.37 * 0.899 = 7 Rn ASRl(2j = 0.174 * 88.2985 * 0.206 ''''''
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M« C T ^ n . ^ T- MCDT ASRT(J) * W T F D ( j - l ) , , , . Mean Solid Residence Time = MSRT = sUM OF"WTFD—^ * J = 11, 2
ASRT(n) * 4.0 ^ ASRT(IO) * 8.0 . ASRT(9) * 13.0 ., ASRT(8) * 16.0 191 191 '191 191
X ASRT(7) * 28.0 ^ ASRT(6) * 28.0 ^ ASRT(5) * 23.0 ^ ASRT(4) * 23.0 • "—T91 ^—T91 •"—T91 " • — W r
+ ASRT(3) * 34.0 . ASRT(2) * 14.0 . ^ 191 191 • ^''^
a = From Table 4
b = From experimental value listed in Table 11
SUMWTO: Cumulative wt. ratio of feed
SUMWTl: Cumulative wt. ratio of char
SIZE: Opening size of screen