Ind. Eng. C h e m . R e s . 1987,26, 221-228
22 1
Parameters Governing Biomass Gasification Ali Tabatabaie-Raissi* Department of Mechanical Engineering, University of Hawaii a t Manoa, Honolulu, Hawaii 96822
George J. Trezek Department of Mechanical Engineering, University o f California, Berkeley, California 94720
Several types of tar-free and tar-emitting biomass fuels were gasified in a 0.864-m-high by 0.578m-inside diameter downdraft gas generator. These fuels included charcoal, wood chips, densified refused derived fuel (d-RDF), and densified digested sewage sludge (d-DSS). An instrumentation system was developed and used t o measure the species gas concentration and the profiles of temperature within the reactor. T h e time-dependent gasification rates for the various fuels were also determined. Due t o batch operation, the principal periods corresponding t o devolatilization and char gasification were observed. Results indicate that the gas quality during these periods cannot be optimized in a single reactor. Furthermore, the initial type of fuel is unimportant during the char gasification period. This conclusion is supported by the results of the gas composition obtained from a simple analytical model for the char gasification region in which water-gas shift equilibria is assumed t o proceed. An example illustrating the use of the model and experimental results for designing and predicting the gasification parameters for similar gas generators is given. The principal objective in gasification is to produce a mixture of combustible gases from a carbon-rich solid fuel. The energy content of the synthesis gas will depend on a number of factors, such as reactor type, fuel type and form, method of heat transfer to the solid particles, ash type, and reactor pressure, to name a few. A downdraft unit was made to evaluate the parameters governing biomass gasification. Typically this mode of operation yields a producer gas with a relatively low amount of tar (less than 10% of that produced in updraft gasifiers). The advantage of using a downdraft gas generator is its low sensitivity to slag formation, charcoal dust, and tar content of the solid fuel as well as its adaptability to load changes (high turn-down ratio). The typical method of operation is to bring air through a central nozzle (tuyeres) or through a number of nozzles in an annular arrangement. The combustion zone is formed near the nozzles, and from there the gases move downward through the gasification zone. With fuels containing large amounts of volatile matter and tar,a pyrolysis region develops near the combustion zone. Gases, tar, and water vapor formed in the pyrolysis zone pass through the bed of hot char beneath, where they crack and form simpler molecules and char. An important result of this reduction is called “flame stabilization” in which the temperature is maintained within a range from 800 to 1000 “C (Desrosiers, 1979). The fuels used for gas generation may be classified into two main categories: (1) “tar-free”, i.e., prime grade charcoal and coke; and (2) “tar-emitting”, i.e., wood, peat, MSW, and to some extent, incompletely charred charcoal. In general, the ash content of tar-free fuels is higher. The factors which affect the design and operation of the gasifier are the chemical composition and configuration of the fuel and the subsequent product ash. The objectives of this study are to develop an insight and quantify, to the extent possible, the effect of the parameters controlling the gasification of biomass. These parameters included the type of fuel in terms of its physical and chemical properties and configuration and the rate of gasification.
Analytical Approach The dominant reactions in the gasification zone of the downdraft reactor are assumed to be (Amundson and Arri, 1978; Yoon et al., 1978; Kosky and Floess, 1980; Biba et al., 1978) 0888-588518712626-0221$01.50/0
C
+ H20
CO + Hz co*2 2 c o i -
(1)
c+ C + 2H2 2 CH4 CO + H20 F? CO2 + H2
(2) (3) (4) Primary reactions in the reaction control volume, Figure 1, are assumed to be
c + 1 / 2 0 2 = co
(5)
Additional assumptions made in developing the model are as follows: (1)the gasification and combustion zones are combined; (2) combustion utilizes all feed oxygen; (3) the product gases which emerge from combustion-gasification control volume are in water-gas shift equilibrium (eq 4) at the exit temperature from the gasification zone; (4)the reaction with hydrogen (eq 3) is not considered because of the slow rate of methanation reaction; (5) a pseudo-steady-state condition prevails within the control volume. The shift reaction 4 is catalyzed by ash or mineral particles of the fuel so that the thermodynamic equilibrium exists at high reaction zone temperatures, T,, i.e.,
The equilibrium constant is given by (Yoon et al., 1978; Kosky and Floess, 1980)
K ( T r )= 0.0265 exp
[E]
(9)
Mass balances for carbon, hydrogen, and oxygen are expressed as (10) M c = Neo + Nco2
+ MH2 + 2M0, = J??H20 + 2 f i C 0 2 + N c O
NH,
I\joH20
+ 2@02
+ NH,O
=
P H 2 0
(11)
(12)
The overall energy equation can be expressed as ( N H 2 - M H J A H , - NCo2a~, -Ncoms = EPi[hi(Tr)- &(Ti)]+ Q, (13) 0 1987 American Chemical Society
222 Ind. Eng. Chem. Res. Vol. 26, No. 2, 1987 HEAT LOSS I 0 ,
r---------I
FLOW IN
I
A I R COMPRESSOR
REACTION
PRODUCT GASES OUT
FLOW M E T E R
t
FUEL IN TEMPERATURE PROBE
Figure 1. Control volume for mass and energy balance.
The solution of the above set of equations is affected by rearrangement into the following expressions, i.e., = [(
NH,o
pi&i)/(l
i=O,,N,,H,O
+
- el) - I \ j o H 2 0 m ,
+
+ (2Mc - p ~ - ~ o 2 N o o , ) ~ ~ ]- /m (j ~ -~my)(14)
( ~ H ~ 2N00, o - &fc)m6
and [(NH?
+ 2WOz + 2M0, - MC - N H 2 0 ) + M H , - NH,?)] / [(2&C - p H 2 0
(NOH&
2p02
- 2M02 -
+ NH~O)N~H,O] = 0.0265e7s60:RT7(15)
where tl = Ql/QT. By use of the average molecule CH,O, for gasified char during the period of quasi-steady-state. gasification, the respective relations for Mc, M H , , and Mo, are Mc = g G / M p (16) a a . UH,= -2( g ~ / h ? ~=) ~ M C (17)
P
Uo, = -@G/&F) 2
P .
=~ M C
(18)
Similarly,
No, = 0.21ga/Ma
(19)
N N 2= 0.79g,/Ma (20) The distribution of dry product gas from the gasification-combustion zone of the reactor could then be determined after N H 2 0 and T, were obtained from eq 14 and 15.
Methods and Materials The details of the design, instrumentation, and measurement techniques and the experimental protocol are considered in this section. On the basis of a design by Cruz (1980), a 0.864-m-high and 0.578-m-inside diameter, downdraft batch-fed gasifier was fabricated. A schematic diagram of the gas generator and the experimental system setup is shown in Figure 2. The unit contained a 0.128-m2, performated stainless steel grate with either 0.006-m-dim e t e r or 0.013-m-diameter holes, an internal 0.08-m-thick insulating layer, and a 0.164-m-i.d. opening, feeding port. The air distributor consisted of a 0.164-m-o.d. and 0.16m-tall steel pipe, closed at one end, with 12,0.003-m-i.d. air tuyeres (4 holes at each section). Producer gas was dischared through a 0.05-m-diameter pipe a t the bottom of the gasifier. To separate particulates and condensables, the gas passed through an aluminum chamber, fin heat exchanger, and a filter before it was discharged into the atmosphere. A 5-hp compressor was used to provide input air into the gasifier. The air flow rate was measured on a 0.472-scms (scms = standard cubic meter per second) flow meter, accurate to 2% of full scale. The gasification rate was measured on a 900-kg Toledo balance (with A0.25-kg accuracy) on which the whole unit was placed. Special probes (Tabatabaie-Raissi, 1982) were designed and fabricated to carry out the task of simultaneous temperature measurement and gas sampling in the reaction
BALANCE
Figure 2. Experimental system setup.
zone of the reactor. Sampling ports were located at 0.15, 0.2, and 0.41 m from the top. The fourth sampling port was located 0.016 m below the air distributor, with a probe extending about 0.06 m from the gasifiers’ interior wall into the reaction zone. The reactor was operated at positive pressures slightly above atmospheric (15-80 kPa, gauge). When tarry fuels were used as feed stock (such as wood chips, densified refuse derived fuel (d-RDF), and densified digested sewage sludge (d-DSS)),the gas sampling lines were purged occasionally by using helium to remove possible condensate buildup. A gas cleansing system, consisting primarily of a heatexchanger arrangement, was used to separate and collect condensables (tar and water vapor) from effluent gas exiting the reactor. Following each run, the condensate was removed, and at room temperature, the total amount of tar generated was determined. The uncertainty regarding the total tar determination was estimated to be *lo% (weight basis). Gas samples were drawn from different locations in the reaction zone of the gasifier and the off-gas line by using a specially fabricated gas sampling system as shown in Figure 2. Each sample was immediately analyzed by using a Hewlett-Packard gas chromatograph (Model 5840A) to give concentrations of the constituent gases. Since gas samples taken during the experimental runs contained essentially no water vapor, the results of the gas analysis were determined on a dry basis. To correct for inaccuracies occurring as air enters the gas during or after the sampling, a procedure (Tabatabaie-Raissi, 1982) was used to convert the results of gas analysis into an “air-free” basis. The axial temperature profiles were determined with 10 Type K (Chromel-Alumel) thermocouples (TC) installed along a 0.65-m-long and 0.013-m-0.d. stainless steel tube every 0.07 m from the tip. The Omegaclad stainless steel sheathed thermocouple wires were used to connect shielded TC junctions through the protective tube to the outside terminals. Similarly, the radial temperature profiles within the gasifier were measured at five locations every 0.051 m between the reactor interior wall and the central axis by
Ind. Eng. Chem. Res. Vol. 26, No. 2, 1987 223 Table I. Data for Fuels and Some Byproducts of Experimental Gasification Runs 3 9 4, 5 7 run no. fuel charcoal charcoal scrap wood redwood apparent density of fuel, pa, kg/m3 440 N/M" 810 N/M bed porosity, tb 0.55 0.52 0.43 0.585 ultimate analysis 57.8 61.11 47.5 50.98 C, 70 by weight 2.13 1.95 5.81 5.83 H, 70 N, 70 1.02 1.34 0.09 0.2 0, 70 N/M 46.1 N/M 18.8 residue N/M 0.5 20.25 others 0.3 ash, % 22.75 22 N/M moisture, % 5.4 6.1 10.4 19 higher heating value, dry basis, MJ/kg 22.3 22.47 20.57 20.68 a
N / M = not measured.
7 wood tar
8 d-RDF 1150 0.491
8 d-RDF tar
10 d-DSS 1360 0.62
65.76 6.88 0.56 25.59 1.0
41.42 5.75 1.55 13.02 37.8
73.57 10.42 0.84 14.01 0.9
30.76 5.15 5.87 24.02 34.2
neg.*
9.5 8.5 18.23
neg.b
41.6 7.6 15
26.42
29.04
neg. = negligible.
Table 11. Organization of the ExDerimental Runs
run
3 4 5 6
feedstock charcoal briquettes mixed scrap wood mixed scrap wood mixed scrap wood
Lam 0.041 0.030 0.033 0.032
7
redwood chips
0.0255
8
d-RDF
9
charcoal briquettes
0.0127b 0.01-0.05' 0.041
10
pelletized sludge (d-DSS)
0.0254b 0.025-0.1'
air flow rate, kalh 2.56 3.66 1.28 1.90 3.37 4.83 5.23 5.31 2.53 3.92 4.28 3.51 6.41 4.52 8.16 12.59 3.84 9.52 15.86
no. of 1oa d in g s
~ ~ l l / ( d p ) hKunii , i , and Levenspiel (1969). *Meter outside diameter pellets.
using a 0.012-m-long and 0.0064-m-0.d. stainless steel probe containing five Type K thermocouples. Additional thermocoupled (Type T) were installed in the input air and gas outlet lines and every 0.21 m from the bottom of the reactor on the outer metal shell. Temperatures were recorded by a Kaye Instrument, Digistrip 11, digital multipoint recorder with accuracy of (f0.03% reading + 0.3 "C). The thermocouples used were Omega Instruments Types K and T with specified limits of error of f2.2 "C or f0.75% (within a temperature range of 0-1250 " C ) and f l "C or fo.75% (within a temperature range of 0-350 " C ) , respectively. Analysis of the fuel and remaining char for moisture and ash content was done by using a standard laboratory drying oven and a high-resolution Mettler balance, accurate to 0.2 mg. Measurements of the properties and elemental analysis for percent carbon, hydrogen, nitrogen as well as the heating values of the solid fuels and remaining chars are summarized in Table I. The analysis of control samples of pure Avicel cellulose provided an estimate of the expected margins of error equal to f2.570 and f l % in determining carbon and hydrogen fractions, respectively. The downdraft gasifier was batch operated, and fuel was fed from the top periodically. The reactor was started with self-igniting charcoal, spread uniformly over the grate. Once a uniformly glowing red charcoal bed was observed, the fuel was loaded and the air flow to the gasifier was started. The upper limit for the air flow rate was set experimentally by the maximum temperatures attained in the reaction zone, which resulted in ash fusion over the
grate location from bottom, m 0.31
grate pinhole size, cm 0.635
air distributor nozzles all holes open
0.31 0.178
1.27 1.27
all holes open all holes open
0.178
2.54
all holes open
0.178
2.54
all holes open
0.178
2.54
0.178
2.54
bottom holes blocked bottom holes blocked
0.178
2.54
bottom holes blocked
Meters long.
grate. This caused partial clogging of the gas passages to occur, increasing the internal pressure of the reactor beyond a safe limit. The lower limit for the air flow rate was determined by the minimum reactor temperatures required to maintain an efficient reduction of the fuel. Summary of the operating conditions is given in Table 11. Temperatures were recorded at 5-15 min intervals immediately after system startup. Gas samples were taken and analyzed every 20 min. Prior to each experiment, the average fuel particle size was determined from representative samples of 30-50, randomly selected, individual chips or pellets, according to a method suggested by Kunii and Levenspiel (1969). At the end of each run, the residual char and ash were collected and weighed. The grate grid opening was slightly changed and modified in the course of the experiments in order to facilitate the removal of ash and residual char as indicated in Table 11. To protect the grate, its location was experimentally chosen to be 0.178 m from the bottom, sufficiently away from the hot reaction zone of the reactor. The temperature data, gas compositions at various locations, gasifier weight loss, and gasification rates were recorded for later reduction by means of a digitial computer. The higher heating value (HHV) of the gas was calculated based on the relationship (HHV)drypas =0.3985+ ~ ~0.1264Xco ~~ + 0.1277~H, (21)
224 Ind. Eng. Chem. Res. Vol. 26, No. 2, 1987
Eool 700
0 Y
K
3 I-
a
K Y
a
f
I-
v)
a
CI
Figure 3. Experimental profiles of temperature and gasifier weight loss vs. fuel residence time in the reactor, d, = 26 mm redwood chips.
where nitrogen in the fuel is assumed not to participate in the reactions. Since the molar rate of inert nitrogen is constant for a given inlet air flow rate, the mass flow rate of the dry producer gas could be evaluated from
Combining eq 22 and 23 will yield 100PN,h;rpC RCN = g a + RG -
(24)
XN,
Equations 23 and 24 could be used to reduce the mass flow rates of dry gas and condensables a t any instant of time from the dry gas analysis and the measured gasification rates. These equations could be integrated with respect to time, to determine the total mass of the dry gas and condensables produced during each experimental run. An estimate of the magnitude of free convective heat loss from the reactor shell was also made (Holman, 1976).
Discussion of Results Due to batch operation, the behavior of the fuel during a gasification cycle, i.e., through the dryout, devolatilization, and reduction periods, could best be analyzed by considering the thermal behavior of the gasifier. Figure 3 depicts the typical temperature profiles within the reactor. In general, three operating time zones were identified during a typical gasification cycle. 1. There was a 5-10-min start-up period from the moment the reactor was loaded until the air flow to the gasifier was started. During this period the gases within the reactor were partially cooled as indicated by the apparent drop in the temperature profiles. 2. There was a transient regime where gas temperatures in the reactor gradually increased due to the heat transfer from the reaction zone. Here the combustion zone shifted from just above the grate (where the fuel was originally ignited) up to a region next to the air distributor holes. The gasifier weight loss was also much more rapid due to drying and partial devolatilization of the fuel. A substantial portion of the volatile matter of the fuel was normally gasified by the end of this period. A portion of water vapor which is generated originates from the breakdown of the tarry fuel in the reaction zone of the gasifier. 3. There was a relatively constant rate of gasification and pseudo-steady-state gas temperature profiles. During this period, the process of solid devolatilization is completed, and from then on the gasification of the remaining char prevails. The temperature profiles and various zones within the reactor become well established.
I
I on
,o
t
loz
103
104
10 5
10 6
AXIAL DISTANCE FROM THE TOP OF G A S I F I E R , m
Figure 4. Pseudo-steady-state axial temperature profiles in the gasifier for various feedstocks and air flow rates.
The pseudo-steady-state axial gas temperature profiles in the reactor for several different types of biomass chars and input air flow rates are given in Figure 4. The location of the maximum measured gas temperature in the reactor remained close to the bottom holes of the air distributor for runs made with wood chips, while it was located near the middle air nozzles for runs with charcoal briquettes and d-RDF pellets. This is due to the fact that the bottom air holes on the air distributor were closed during the later runs, to ensure that the combustion zone in the reactor remained within the constricted throat area of the gasifier. The results of Figure 4 indicate that the combustion zone temperature should increase as the air flow rate is increased for each type of biomass char. The maximum gas temperature at the combustion zone is a strong function of such parameters as the air flow rate and temperature, moisture content of the fuel and input air, extent of gas-phase combustion, particle size, composition of ash, and thermal conductivity of the fuel bed. The reactivity of fuel is especially important as it correlates inversely with maximum gas temperatures attainable in the reactor due to the endothermicity of the gasification reactions. For the experimental conditions of this study, wood chars appear to be more reactive than d-RDF chars and charcoal (with which reaction zone gas temperatures in excess of 1000 "C were recorded, Figure 4). The reactor gas temperature gradients in the fuel bed above the combustion zone appeared to be relatively small for most cases. The temperature of the gas, however, dropped sharply through the void region above the fuel bed (Figure 4). Due to the bulk mass transfer of the reaction products toward the bottom of the reactor, the axial gas temperature gradients were greater below the air inlet nozzles. The measured gas temperatures in radial and tangential directions within the reaction zone indicated a nearly uniform radial temperature profile to exist across the reduction zone, 0.016 m below the air distributor for all the biomass materials examined. This was especially true for experiments with charcoal briquettes and, in general, during the pseudo-steady-state operating period of the reactor. It was found that there was a considerable temperature variation in the tangential direction in the combustion zone of the reactor. In the reduction zone, however, this dependency was found to be less apparent. In Figure 5 , the percentage of fuel converted is plotted as a function of time for all biomass fuels examined.
Ind. Eng. Chem. Res. Vol. 26, No. 2, 1987 225
0
4c
e€
120
180
TIME, ~
m
240
280
320
~ r ~ ~ i . . i
Figure 5. Percent fuel conversion vs. gasification time.
Generally, higher fuel conversions were obtained as air flow rates to the gasifier were increased (at constant fuel residence times). Furthermore, a t a fixed input air flow rate, the amount of charcoal converted was found to be approximately 1 order of magnitude less than that of the other fuels. Once the fuel was completely devolatilized, the rate of weight loss (conversion) declined substantially. The rate of the gasification of the remaining char cannot be improved without a substantial increase in the input air flow rate which in turn elevates the temperatures in the reactor and reduces the quality (calorific value) of the gas. Figure 5 also provides information regarding the reactor loading frequency. The gasification rate is strongly affected by the physical and chemical properties of the fuel. For tarry biomass fuels with relatively high moisture content (up to 20% in the case of wood), the gasification proceeded with much lower reaction zone temperatures and provided a gas with relatively high calorific content (as high as 9 MJ/scm). The gas also contained some concentration of tar (as much as 0.065 kg/kg of ash-free for wood and d-RDF). On the other hand, charcoal and d-DSS required relatively high input air flow rates, yielding a gas with comparatively lower heating values. The total dry gas yield from gasification of tarry fuels such as wood and d-RDF was found to be much lower than that of charcoal and d-DSS. It is important to realize that the gasification rate changed once the fuel was partially devolatilized. From then on, the gasification process proceeded at a much slower rate (given a constant air flow rate) and resembled that of the char gasification processes. The calorific content of the producer gas also declined, requiring the addition of fresh fuel and removal of the remaining char from the reactor. This then determined the frequency with which the reactor had to be loaded to ensure gas production with an acceptable energy content. Since the devolatilization occurred relatively fast (followed by a long period of char gasification), the initial high quality of the gas was difficult to maintain over the entire period of operation. For example, the extent of gasification of the d-RDF pellets should be in the neighborhood of 50%, consistent with the flat portion of the corresponding curves in Figure 5. This can also be seen in Figure 6 where the total quantity of dry gas produced during gasification is given in terms of percent fuel conversion. Note that from Figure 6, in order to avoid char gasification, the extent of d-RDF converted should be kept in the neighborhood of 50% or less. The total dry gas yield appears to be a function of the volatile matter content of the fuel. In general, the data indicates that this yield is not strongly influenced to any large extent by the input air flow rates used to gasify the fuel. Data points for the gasification of dry tar-free charcoal at air flow rates of 8.16 and 12.59 kg/h fall in between the limiting lines for complete com-
= CO+1.88h
/
/*
%GWL,percentweight loss, kg / k g l s h , , ~
Figure 6. Total gas yield as a function of percent fuel conversion for several types of biomass fuels under various conditions. I
T I M E , mlnutea
Figure 7. Producer gas and combustion zone, gaseous concentrations, and higher heating values for gasification of 12-mm-0.d. d-RDF pellets a t gdrran= 3.51 kg/h.
bustion and gasification of pure carbon in the air as shown in Figure 6. The total yield of dry gas from gasification of d-DSS appears to be much higher than that of wood or d-RDF. This is due to the low conversion of volatile matter during the d-DSS gasification. A t about 50-607'0 conversion, the yield begins to increase sharply for wood and d-RDF. The total amount of dry gas produced from gasification of charcoal briquettes is found to be more than twice as much as that from d-RDF within the entire range of the conversion. This is apparently due to much higher input air flow rates and reaction zone temperatures required for gasification of tar-free materials. Figure 7 depicts the trade-off which exists between the heating value of the producer gas and the conversion efficiency of the process for gasification of 0.013-m-0.d. dRDF pellets at an air mass flow rate of 3.51 kg/h. The thermolysis of tarry fuels results in pyrolytic gases, tar, and a carbon-rich residue (char). Emission of this pyrolysis gas during devolatilization of the fuel is responsible for enriching components of the gas, such as methane and
226
Ind. Eng. Chem. Res. Vol. 26, No. 2, 1987
z . 2
w J
BO
?
'5
110
'15
120
w
E I u)
& I-
I
Y
5
c z 2 K
w
n 0 2
w> W
t
GAS CONCENTRATION
v)
z
xz
8
ICaISuI~I.dl
Figure 9. Comparison between experimental and calculated values of the producer gas composition.
2
3
Figure 8. Total yield of condensables VS. percent fuel conversion.
other hydrocarbon gases. Since methane is known to be the only pyrolysis product which is never completely converted in the reduction zonc (Groeneveld and Van Swaaij, 1980), an estimate of the extent of fuel devolatilization can be made by means of concentration of the methane in the product gases. As the solid fuel gradually decomposed and gasified, the concentration of methane in the product gas was sharply decreased. A decline in the quality of the gas and its higher heating value was also observed as the conversion continued (see Figure 7). This is apparently caused by complete devolatilization of the biomess and the prevailing char gasification processes. The general trend of the data indicates that during gasification, the gas heating value reaches a maximum at a fuel conversion of 10-60% depending upon the type of biomass gasified. For example, for d-RDF, examined in this work, the maximum appears to be a t approximately 25% conversion (see Figures 5 and 7). Beyond this limit the quality of the gas in the reduction zone of the gasifier began to decline. It is also important to note that once the fuel was completely devolatilized, hydrogen almost disappeared from the constituent gases emerging from the gasifier. The air flow rate then had to be increased in order to improve the gas quality, as discussed previously. The total yield of condensables (tar + water) for several types of biomass have been compared in Figure 8. The total amount of tar generated during the experimental runs with wood and d-RDF was found to be approximately one-tenth to oneeighth of the total condensables collected. These results are consistent with those of other investigators (Groeneveld and Van Swaaij, 1980) indicative of the fact that tar evolves in the product gas due to either incomplete conversion in the combustion zone or incompletely devolatilized char entering the reduction zone of the reactor. The values of the gas composition obtained from the analytical model developed for the char gasification period of operation are compared to the experimental values in Figure 9. The input data required for the simulated gasification runs and computation of the output conditions are given in Table I11 for several biomass fuels used. The results indicate that, once the devolatilization period has been completed, the influence of the type of biomass char
Table 111. Values of the Input Parameters for Several Simulated Biomass Gasification System Models fuel d-RDF charcoal d-DSS
gG,
NHzO~
kg/h
kg-moli h 0.05 0.03 0.03
1.42 1.13 2
a 0.8 0.3 0.4
P 0.5 0.05 0
Table IV. Sensitivity of the Model Predictions to the Changes of Several Input Parameters (Charcoal Run, g , = 4.52 kg/h) re1 change of values (%) from a 20% increase in the value of the input parameter model input %Np %COP %CO %H2 T,, "C parameters T,, "C neg. neg. neg. -2.9 1.4 -0.14 0.56 -1.8 13 -3.2 NH20 -0.6 2.8 -7.4 17.6 -4.5 MH, -0.3 -2.8 4.9 8.8 neg.
on the product gas composition is negligible. Further, this comparison tends to justify the assumption of homogeneous water-gas shift equilibria. In Figure 9, the results of previous analysis and that of equilibrium have been compared to experimental data corresponding to charcoal. Equilibrium calculations have been performed for a mixture of carbon and air (with 3% moisture content). In these calculations, the charcoal was assumed to consist of only carbon. This might not be completely justified under real gasification conditions, even though the results seem to indicate that the final gas composition from completely devolatilized charcoal resembles that of equilibrium reasonably well. The sensitivity of the model predictions to the changes in the assumed values of several input parameters for a simulated gasification run with charcoal is given in Table IV. This table includes the changes in computed values of the producer gas composition when the input values were increased by 20% relative to the original assumed values of these parameters. Finally, it is instructive to demonstrate, in an example, the way one might use these experimental findings for design and performance evaluation of a similar gas generator. Let us consider the case of a small downdraft gasifier which utilizes 0.013-m-diameter d-RDF pellets to produce a low calorific gas, accommodating a mass flow rate of air equal to 5 kg/h. The proposed design questions are (1) the loading frequency and capacity of the gas generator,
Ind. Eng. Chem. Res. Vol. 26, No. 2, 1987 227 and (2) the amount of the gas produced. (1) For d-RDF using Table I, assume an apparent density of 1150 kg/m3, a fixed bed porosity of q, N 0.49, and an ash content of 10%. A fuel conversion of about 50% is assumed so that
viduals during the course of this study. Special thanks are due to George J. Savage of Cal Recovery Inc., Benjamin E. Gordon and Wallace R. Erwin of the University of California, Laboratory of Chemical Biodynamics, and the staff of J. T. Thorpe and Son, Inc.
f ~ w L= %GWL/100 = 50/100 = 0.5 An estimate of the loading frequency of the gasifier can be made by using Figure 5 and interpolating between runs 8-1 and 8-2 for an input air flow rate of 5 kg/h and to N (70 + 150)/2 = 110 min. Then GWL = fGWL(1 - Xash)Pa(l - tb)V
(25)
But
From Figure 6, for d-RDF, we have
I,
N
0.45
Similarly from Figure 8
I,
0.38
-
The total mass of the air required is ma = gato N (5)(110/60)
9.2 kg
Then
V=
9.2 (1 - 0.1)(1150)(1 - 0.49)(0.45
+ 0.38 - 0.50) N
0.053 m3
(2) The average rate of gasification of the fuel is estimated from gG
= GWL/to = f G W L ( 1 - Xash)Pa(l - Cb)V/tO gG
= 7.58 kg/h
Nomenclature
d = mean particle diameter, m (&,)h,l = hydraulic diameter of the ith sampled particle, m f~~~= %GWL/100 = fraction fuel converted, kg/kg of ashfree fuel ga = mass flow rate of the oxidant (air), kg/h gCN= mass flow rate of the condensables, kg/h gg = total mass flow rate of the condensables and noncondensables out of the reactor, kg/h g~ = gasification rate of the solid fuel, kg/h, which is equal to the rate at which reactor weight loss occurs (GWLlAt) or rate of fuel conversion g P G = dry producer gas mass flow rate, kg/h gtar = mass flow rate of the tar in the off gases, kg/h gw = mass flow rate of the water in the product gas, kg/h GWL = gasifier weight loss, kg, measured on the scale %GWL = percent fuel conversion, kg/kg of ash-free fuel h,(T)= molar enthalpy of gas i at temperature T , kJ/kg-mol (HHV)d, gas = dry gas higher heating value, MJ/scm Il = totar producer gas yield per unit weight of ash-free fuel, estimated from Figure 6 I , = total condensate yield per unit weight of ash-free fuel, estimated from Figure 8 ma = mass of air, kg Ma= molecular weight of the air = 28.975 kg/kg-mol MF= molecular weight of the gasified fuel, kg/kg-mol M I = molecular weight of gas species i, i E H,, N,,CO, C02, and CH,, kg/ kg-mol M, = molar feed rate of i in the fuel, i H,, O,, and C MPG = Cx,M,, molecular weight of the dry producer gas, kg/ kg-mol of PG y = number of the sample particles yL= molar feed rate of i, i CO, COz, H,, and H 2 0 NO,= molar feed rate of blast gases at air distributor location NoN2= 0.79ga/Ma,molar feed rate of the nitrogen gas in the input air, mol/h QT = total heat released at the reaction zone due to exothermic combustion reactions, kW t = elapsed time, h t o = time period between two successive loading of gasifier, h T, = temperature at inlet conditions, K T , = reaction zone temperature, K V = gasifier capacity, m3 r a s h = ash content of the solid fuel xNS,x , = percentage of nitrogen and gas species i in the dry product gas Greek Symbols
Then gg = g G
+ ga = 7.58 + 5 = 12.58 kg/h
An estimate of the quality of the gas produced could also be made by using an average molecule CH1.6700.24N0.032 for d-RDF from Table I, with an assumed value of g, and calculated value of g G . An average value for the moisture content of the fuel and air must be assumed in order to calculate the composition of the producer gas and its higher heating value.
Acknowledgment The financial support for this work was provided by the US Department of Energy Contract DE-AT03-79ER15391 and US Environmental Protection Agency under Grant R-807011010, which is gratefully appreciated. We acknowledge the support and contribution of many indi-
CY,
p = moles of hydrogen and oxygen per 1 mol of carbon
AH, = molar heat of reaction j as written, kJ/kg-mol tb tl
= fixed bed porosity = ratio of reactor heat loss to total heat released in the
reaction zone of the reactor due to exothermic combustion reactions pa = apparent density of the solid fuel x P G = summation over permanent gases of hydrogen, nitrogen, carbon monoxide, carbon dioxide, and methane Literature Cited Amundson, N. R.; Arri, L. E. AIChE J . 1978, 24, 87. Biba, V.; Macok, J.; Klose, E.; Malecha, J. Ind. Eng. Chem. Process
Des. Deu. 1978, 17, 92. Cruz, I. E. In Thermal Conversion of Solid Wastes and Biomass; Jones, J. L., Radding, S. B., Eds.; American Chemical Society: Washington, DC, 1980; pp 649-670.
228
Ind. Eng. Chem. Res. 1987, 26, 228-236
Desrosiers, R. In A Survey of Biomass Gasification; Reed, T . B., Ed.; National Technical Information Service: Washington, DC, 1979; Vol. 11, SERI/TR-33-239. Groeneveld, M. J.; Van Swaaij, W. P. M. ACS Symp. Ser. 1980,33, 130. Holman, J. P. Heat Transfer, 4th ed.; McGraw-Hill: New York, 1976. Kosky, P. G.; Floess, J. K. Ind. Eng. Chem. Process Des. Deu. 1980, 19. ,586.
Kunii, D.; Levenspiel, 0. Fluidization Engineering; Wiley: New York, 1969; pp 61-69. Tabatabaie-Raissi, A. PbD. Dissertation, University of California at Berkeley, Berkeley, CA, 1982. Yoon, H.; Wei, J.; Denn, M. M. AZChE J . 1978, 24, 885.
Receiued for review October 10, 1983 Revised manuscript receiued June 11, 1986 Accepted July 17, 1986
Mass-Transfer Characteristics of Valve Trays Richard D. Scheffe and Ralph H. Weiland* Department o f Chemical Engineering, Clarkson Uniuersity, Potsdam, N e w York 13676
Gas- and liquid-side mass-transfer coefficients and gas-liquid interfacial areas for Glitsch V-1 valve trays are reported. T h e data were taken from a three-tray column, each tray having an active deck area of 0.372 m2 at a tray spacing of 0.61 m. Gas and liquid rates covered the range from severe weeping to loading conditions, and weir heights were varied from 22 to 150 mm. The results are presented in terms of Sherwood numbers for mass transfer as functions of gas and liquid Reynolds numbers, dimensionless weir height, and Schmidt numbers. In the design of tray columns for separation processes, the usual procedure is to determine the number of equilibrium stages required and then to apply a correction for stage efficiency. To compute these stage efficiencies, one needs information on the number of transfer units for the gas and liquid phases or, equivalently, the gas and liquid side mass-transfer coefficients. For processes involving liquid-phase chemical reactions, on the other hand, the equilibrium stage model corrected for efficiencies is not a workable approach at all. This is because of the difficulty in computing the highly composition-dependent tray efficiencies in the first place, as well as the fact that efficiencies can vary quite markedly from tray to tray and component to component. In such cases, column design and simulation is more sensibly approached from the standpoint of a mass-transfer rate process. Liquid-side transfer coefficients are corrected for chemical reaction effects through the use of enhancement factors, and the whole efficiency question is circumvented altogether. This formulation often requires interfacial areas to be known separately from the individual-phase mass-transfer coefficients. Although a fair amount of mass-transfer data are available for various packings (Au-Yeungand Ponter, 1983; Bravo and Fair, 1982; Linek et al., 1984), similar information for various types of trays is quite sparse. The literature up to 1966 has been nicely summarized by K&tAnek and Standart (1966). Information on bubble-cap trays is the most complete (AIChE Bubble Tray Design Manual, 1958), and further data appearing since 1966 include those of Sharma et al. (1969), McLachlan and Danckwerts (1972), Burgess and Calderbank (1969) and Kochetov and Rodionov (1977). These latter works, however, added little to the available data base. Mass-transfer-coefficient data for sieve trays (Calderbank, 1959; Calderbank and Moo-Young, 1961; Harris and Roper, 1963; Sharma and Gupta, 1967; Pasiak-Bronikowska, 1969; Pohorecki, 1968; Counce and Perona, 1979) tends to be much more scanty due to the wide variation in tray designs, the fact that sieve trays can be operated both with and without downcomers, and the use in some cases of a tray with a single sieve hole. To our knowledge, there are no mass-transfer data based on gas absorption
processes available in the open literature for the more modern valve or ballast tray; there are, however, voluminous tray efficiency data based on distillation experiments which is some instances can be used to calculate gas-side mass-transfer coefficients. Here we report gas- and liquid-side mass-transfer coefficients and effective interfacial areas for V-1 type valve trays manufactured by Glitsch, Inc. The data were taken on a three-tray column, each tray having an active deck area of 0.372 m2 at a tray spacing of 0.61 m. The central tray was the test tray, and measurements were made by using the now-standard procedures of chemically reactive absorption. Flow conditions through the column were varied from severe weeping to entrainment conditions, and weir heights were varied from 22 to 155 mm. The results are presented in terms of Sherwood numbers (dimensionless mass-transfer Coefficients) and interfacial areas as functions of gas and liquid Reynolds numbers, dimensionless weir height, and Schmidt numbers, and so they should be able to be used for any system of known physical and transport properties.
Experimental Section Equipment. Schematic diagrams showing the experimental setup and the principal dimensions of the column are shown in Figures 1 and 2, respectively. The column was of rectangular cross section and was constructed of 19-mm cast acrylic sheet bolted together and sealed with silicone caulk. The trays were Glitsch V-1 ballast type made of 14-gauge 304 stainless steeL The trays were square, each with a bubbling area of 0.372 m2 on a spacing of 610 mm. Each tray contained 50 standard (1 7/s-in. diameter) Glitsch V-1 ballast units. The column internals were designed with provision for adjustable weir heights from 22 to 155 mm and adjustable downcomer seals. To prevent entrained liquid from leaving the column in the gas stream, a 150-mm thick by 560-mm square demister (Otto York, Inc.) was placed 200 mm above the deck of the top tray. Access to each tray was had through 300-mm by 650-mm doors. The setup permitted liquid circulation rates from about 1.1to 6.9 kg/s (8700-55000 lb/h) and air rates in the range 0.11-1.1 kg/s (200-2000 scfm). All flow rates were mea-
088S-S88S/87/2626-0228~01.~0/0 0 1987 American Chemical Society
