Ind. Eng. Chem. Res. 1988,27, 317-324
317
Modeling an Oil Shale Fluid Bed Combustor Iacovos A. Vasalos,* Alekos Lefkopoulos,+and Maria Georgiadout Chemical Processing Engineering Research Institute and Department of Chemical Engineering, P.O. Box 1951 7, University of Thessaloniki, Thessaloniki, Greece
Oil shale retorting involves heating of solid particles and pyrolysis of the organic matter to produce hydrocarbon liquid shale oil. During the pyrolysis process, part of the organic material remains in the inorganic matrix as coke residue. Combustion of the coke residue can provide the energy necessary for retorting. In this paper the use of a fluid bed combustor to burn the coke residue is examined. The basis for predicting the performance of the fluid bed combustor is the application of the two-phase theory of fluidization. The carbon burning efficiency was calculated as a function of temperature, pressure, and bubble size. For the same conditions, the carbonate decomposition and the associated energy loss were also established. Conditions were found which make feasible complete I carbon combustion with minimum carbonate decomposition. Oil shale retorting involves heating oil shale to a temperature over 480 OC, whereupon kerogen decomposes to form light oil, gases, and carbon residue. A reaction sequence for the retorting process has previously been described (Wallman et al., 1980). Although several processes have been disclosed in the patent literature, this paper will provide useful information to those methods which fit the generalized process scheme in Figure 1 (Vasalos et al., 1984). Figure 1 shows a generalized retorting scheme which represents any process where the retorting reactions take place add in which carbon residue is burned off the retorting solids. Fresh shale particles, generally smaller than 6 mm, are mixed with a hot, inert heat carrier, such as combusted spent shale, in the mixer. Heat is transferred from the hot recycled particles to cold shale particles. Retorting begins in the mixer and continues in the retort. The purpose of the retort is to provide additional time to complete retorting, if needed. The retorted solids are withdrawn from the bottom of the retort and are combusted and pneumatically transported by air in a vertical riser. The combustion of the carbon residue results in a temperature increase along the riser. A t the top of the combustor, the solids are split into two streams, a fine and a coarse fraction. The fine fraction exits the separator with the flue gas, while part of the hot, coarse fraction is recycled to the mixer to provide process heat. Specific examples of processes which have reached an advanced state of development include the Chevron’s Staged Turbulent Bed (STB) process (Tamm, et al., 1982), the Lurgi process (Rammler, 19821, and the Tosco Hydrocarbon Solids process (HSP) (Hall, 1982). Each of the above processes has disclosed variations on both the retort and the combustor section. Since this paper will focus on the combustor section, let us review some of these variations. The riser combustor is considered as the only option in the Lurgi publications, while the Tosco HSP process has emphasized the fluidized bed combustor. Chevron publications have, in addition to the riser combustor, disclosed a modified fluidized bed reactor type similar to the STB concept (Tamm, 1982). In trying to assess the various combustor types, several factors should be accounted for; some of these key factors are listed in Table I. Although this qualitative comparison of the two combustors show no clear champion, it is important to quantify some of these factors before a final selection is made. Special emphasis will be placed in t Current address: Department of Chemical Engineering, University of Illinois at Urbana-Champaign, Urbana, IL 61801.
OSSS-5SS5/SS/2627-Q317~Q~.50/0
predicting the operating temperature, the gas requirements, and the carbonate decomposition. Since no reliable data could be found in the literature, the attrition question is not addressed at this time. This implies that for the application of this model, prior knowledge of attrition data is required.
Mathematical Formulation Factors important in the derivation of the mathematical equations which described the performance of an oil shale fluid bed combustor are the fluidization of the solid particles, the reaction of residual organic carbon with 02,and the decompositionof carbonate minerals. Let us examine these factors before we derive the material and energy balance equations. Fluidization of the Solid Particles. The particle size distribution of the solids contained in an oil shale fluid combustor will depend highly on the origin of the solids. This is shown by data published by Patzer and Carr (1984), and for convenience the data are reproduced in Table V. If the criteria derived by Kunii and Levenspiel (1977) are applied, a system of this type can be represented by the two-phase theory of fluidization. In simple terms, the gas after it leaves the air distributor is split in two portions. Part of the gas is used to fluidize the solids at conditions of the minimum fluidization velocity (umf),while the remaining gas forms bubbles which grow as they rise and displace the solid particles. The initial bubble size depends on the distributor design, while the ultimate bubble size depends on the bed height. The deeper the bed, the larger the bubble diameter, which implies that for deep beds it may be desirable to place internals inside the bed to control the bubble size. Controlling the bubble size to less than a given value makes the oxygen residing inside the bubble more effective because, as it will be discussed later, the gas interchange coefficient is inversely proportional to bubble size. After the bubbles reach the top of the bed, they burst in the surface, causing particles to be injected in the freeboard and some of them to be entrained with the gas out of the combustor. The equations for the various parameters used in the two-phase theory of fluidization are given in Table 11. Reaction of Residual Organic Carbon with O2 (C O 2 COz). Previous investigators (Sohn and Kim, 1980; Braun, 1981) have recognized the importance of oxygen diffusion into the pores of retorted solids to determine the overall rate of carbon burning. For conditions typical in a fluid bed combustor, we use the intrinsic rate of carbon based on the work of Sohn and Kim (1980). The oxygen mass-transfer coefficient from the bulk to the
+
-
1988 American Chemical Society
318 Ind. Eng. Chem. Res., Vol. 27, No. 2, 1988 Table I. Comparison of Potential Carbon Residue Combustors item riser combustor riser exit temp greater than 1200 O F required 1. operating temp
fluid bed combustor flexible due to the increased particle residence time 2. gas requirements determined by the minimum transport velocity of the determined by the carbon combustion particles requirements high due to the high combustion T and residence potentially low due to the lower combustor 3. carbonate decomposition time in the collecting bin temp 4. carbon conversion incomplete due to low residence time potentially high due to flexibility in solids residence time and temp 5. commercialization requirements need to develop scale-up criteria for large size units limited due to the extensive fluid bed coal combustor experience
strong function of temperature. Hence, for the purpose of this study, (1)and (2) were modified by us to reflect the reduction in decomposition rate as the partial pressure of C02 in the emulsion phase, PCo2,,,approaches the equilibrium decomposition pressure for COz from MgCO, and CaC03:
Flue gas fines
+
Products
~~
+ Fines
-)
R2 = KzUz( 1 - Pc02,e
[kg of C02/(m3of shale-s)]
Pc0z9e2 Combustor
withdrawal R3
0 Air Figure 1. Generalized retort scheme.
particle and the effective oxygen interparticle diffusivity are based on data provided in a Lawrence Livermore Laboratory publication (Braun, 1981). The relevant equations for all these effects are reported in Table 111. Item a in Table I11 indicates the explicit dependence of the intrinsic carbon rate on oxygen partial pressure. It is this dependence which explains the increase in combustion efficiency with increasing combustor pressure. Decomposition of Carbonate Minerals. Carbonate minerals are contained in large quantities in the Colorado oil shale. Typical values reported by Baughman (1978) are 32 wt % dolomine and 16 wt % asbestite. It has been stated by Braun (1981) that accurate representation of the kinetics of decomposition is difficult because of the many chemical reactions which can take place in the inorganic matrix. For this reason, we accept Braun’s simplification: “...to consider only two COz containing components. The first component ( U2)represents COz from the MgCO, of dolomite as well as the COz from other relatively lowtemperature decomposing carbonates, such as nahcolite and dawsonite. This is a valid representation only if the latter minerals are present in minor quantities. The other component (U,) represents the C02 from the CaCO, portion of dolomite and from the initial calcite content.” The first-order rate expressions for the decomposition reactions are (Campbell, 1978)
Rz = KzU2 [kg of C02/(m3 of sha1e.s))
(1)
R3 = K3U3 [kg of C02/(m3 of shale-s)]
(2)
(3)
( i,::,r3)
=K3U3 1--
[kg of C02/(m3of shale41
(4) where Pc0,,,, = C 0 2 decomposition partial pressure for MgC03, Pc02,e,= C02 decomposition partial pressure for CaCO,, and Pco2,,= COz partial pressure in the emulsion phase. The equations for the decomposition pressures Pco2,e, and P c o ~ ,are ~, 60863(T - 673) (5) log = 4.571 X 673T
(-)
log Pco2,e3= --11355 - 5.388 log T + 26.238 (6) T where T is in K. Material and Energy Balances. The equations for the oxygen and carbon dioxide variation along the bed height were formulated based on the following assumptions: 1. The emulsion phase is well mixed, which means it has constant temperature and concentration of all species. 2. The bubbles move along the bed in plug flow with concurrent exchange of mass with the emulsion phase. 3. The system operates at steady state. With the above assumptions,the following equations are applicable: oxygen balance in the emulsion phase: H
(l -
- CO,)+ € b A K p L (C02,bcoze)dZ - (1 - €b)AH(1 - emf)R8/32= 0 ( 7 )
tb)AUmf(C02,,n
carbon dioxide balance in the emulsion phase: (1 - €b)AUmf(O - cC02,e)+
and where K 2 = 1.7 X 1O’O exp(-29090/T) [s-’1, H 2 = -3.0 X lo6 [J/kg of C02], K3 = 9.6 X 1O1O exp(-36050/T) [s-l], and H3 = -2.9 x lo6 [J/kg of CO2]. According to (1)and (2), the rate of decomposition is not restricted by the rate of recombination of C02 with MgO and CaO. However, according to data published in the InternationaZ Critical Tables (1930), the partial pressure of C02 resulting from thermal decomposition of the MgCO, and CaCO, is a
C ~ A K ~ L ~ ( C-CCCO~,.) O , , ~dZ + (1 - edAH(1 - emf) X (Rz + R3)/44 (1 - €b)AH(1 - ~,f)R8/32 = 0 (8) oxygen balance in the bubble phase: dCo,,/dZ = Kp(co2e - ‘0Z.b) where COz,b(Z = 0) = Coz,m COz balance in the bubble phase: b‘
ub
dCco,,b/dZ = K p ( C C 0 2 ,
- ‘Cozb)
(9)
(10)
Ind. Eng. Chem. Res., Vol. 27, No. 2, 1988 319 Table 11. Fluidization Parameters of an Oil Shale Fluid Bed Combustor item eauation used
Table 111. Carbon Burning Kinetics item equation used RE,= KJ'02U6/f20 kg Oz/(m3shales) a. Rs. where
K8 = 1.52 exp[ -11092(
where De,02= 1.04 C.
X
-
$-)I
10-15@F,66m2/s
R& = 3KdCo2/rj
R& where
c. distributor design
= 0.3AF'h
AF'd
--)
112
2gcud
uor = 0.8(
7r
0.2952 - 0.4671 In Re
uo = ~do,2U,tNor
0.03366 (In Re)2
Dbo = 0.347[A(u0- U ~ ) / N , , ] ~ ~ '
d. bubble parameters
Re =
Dbm = 0.626A(uO- U,f) Mori and Wen-model 1
- Db
Dbm
Dbm - Dbo Db,WM
2yPG 1.5~(1 e)
= exp(-0.3Z/D)
(Dbm- Dd[exp(-0.3H/D) - 11 0.3(H/D)
= Dbm +
+
Cranfield and Geldart-model 2 Db = 0.0326(u0 - Umf)1'11p'8' 0.0326(u0 - Umf)1'11110.81 Db,CG =
1.81 Cranfield and Geldart ( 5 grids)-model 3
1
1 1
1
R8a
R8b
R8c
-+-+-
kg 0 2 / ( m 3 sha1e.s)
where CCOPb(Z = 0) = 0. It is convenient to obtain an explicit solution to the above set of equations by using an average bubble diameter:
0.0326(u0 - ~ ~ 3 ~ 9 H / 6 ) ~ ' ~ ~
=
Db,CGS
1.81
+ 22.26(Db)'I2
uo - U,f
ub e. eb
6b
=
UO
-U
d
-
ub K, = ll/Db
f. K,
g. U t Ut
Ut =
=
[225
h. K*
A P S - Pg)dpi2
Re, < 0.4
law ("
]
- pg)2g2 pgcl
112
dpi 0.4 < Re, < 500
It is realized that a numerical solution will give greater accuracy. However, it is more important to obtain a more convenient form for engineering applications. The approximate solutions are reported in Table IV. In addition to the material balances on the emulsion and the bubble phase, the final solutions must also satisfy the following overall balances for the combustor: carbon balance:
K* P A
ut/u0 < 1, dpi < 830 pm _.-
Finally, the solution must be consistent with the overall combustor energy balance:
PgUo
ut/uo < 1, 830 p m
K* = 49.1 ex.( PgUo
< d,; < 2000 pm
-4%)
ut/uo > 1
where TI is the reference temperature.
320 Ind. Eng. Chem. Res., Vol. 27, No. 2, 1988 Table IV. Oxygen a n d Carbon Dioxide Solutions
C 0 2 concn in emulsion phase: C C O ~=, ~ [(CC)(AD)(PRMG + PRCA)
+
CC = 1/RT B = (1 - +,)HA(l - emf) T = (1 - t b ) A ~ , r AD = B/(44(T - Y)) Y = cbubA(exp(-N1) - 1) Nl = K&/Ub PRMG = K2U, PRCA = KSU3 COPconcn in the bubble phase:
Table V. Entrance Conditions for Fluid Bed Combustor (Mostly from Patzer and C a r r (1984))" gas phase: temp, K 573 rate, mol/s 2430 solids: 0.8 sphericity density, kg/m3 inert 1682 reactive (3.66 wt % C) 1734 temp, K 798 distribution, wt % class inert reactive total 0.0044 5.8 0.0 5.8 0.0116 8.7 0.0 8.7 0.0258 13.6 9.3 13.6 0.0408 21.7 0.0 21.7 14.9 0.0650 14.9 0.0 0.0900 8.4 1.9 10.3 9.2 5.3 14.5 0.1300 0.2320 2.0 3.5 5.5 4.1 0.3500 1.0 3.1 0.4500 0.0 0.9 0.9 total 85.3 14.7 100 OBasis: loo00 ton/day module; 125 L/ton shale; 12 combustor.
"1 For the calculation of the total amount of carbonates decomposed, it is very important to realize that it is possible that some retorted particles may be recycled in the system many times. This will depend on the shale richness and the degree of decrepitation occurring during combustion. As stated by Hall (19821, "with high grade, mahogany zone Green River shales, the residual ash decrepitates", while "when processing low grade shales, a competent ash normally results which can be used directly as the heat carrier". For this paper, it was assumed that the average shale particles will go around the system N times where N is an integer close to the ratio N = rate of retorted solid entering combustor/ rate of oil shale to retort (16) The total amount of C02 produced from decomposition of MgC03 and CaCO, is then defined by
SFCO,%=
Method of Solution. Since the purpose of this paper is to study the effect of fluidization parameters and operating variables on the percentage of carbonates decomposed (Wed) and the percentage of heat of combustion of carbon, hydrogen, and sulfur absorbed by the endothermic decompositionreactions, a flexible computer program was written. The final solution accounts for all variables listed in the Nomenclature section. Some important considerations include the following: (a) description of the retorted, spent solids and fines by several discrete particle size fractions (as shown in Table V); (b) calculation of all
17 m
X
Co.bustnr 1:922 OK
.mtYMcYRI
150
t ; i
/
s
W S
0
1.5
BED'HEIGHT.
H1'5
7.5
I
Figure 2. Mean bubble size va bed height.
stream properties using literature data (heat capacities, viscosities); ( c ) characterization of the fluid system with all system parameters listed in Table 11; (d) calculation of kinetic parameters according to Tables I11 and (1)-(6); (e) calculation of oxygen and carbon dioxide concentrations at the top of the bed with equations reported in Table Iv; (f) calculation of carbon conversion, carbonate decomposed, and final material and energy balances by applying (12)-(15). It is understood that (7)-(10) can easily to numerically integrated by using standard mathematical )techniques. Since the explicit solution method using an average bubble size gave similar results, it is reported here for easy use in industrial applications.
Results and Discussion The basis for these calculations is the design parameter reported in Table V (Patzer and Carr, 1984). The particle size distribution of the solids entering the fluid bed combustor varies widely, dependent on whether the particles are reactive (retorted solids) or inert (spent solids). In this particular example, the average particle diameter of the reactive solids is 1600 pm and that of inert solids is only 270 pm. Four topics will be discussed for this particular system: the average bubble size based on the theoretical calculations, the variation in oxygen concentration along the bed
Ind. Eng. Chem. Res., Vol. 27, No. 2 , 1988 321
m
2 S
9.6
FDUL BED HEIEN
3 Y
-
Caam0;lSm T:W4 OK
0
wz.3 n
FNIM
BED MIEN w3.6 n
FWL
nm WEN w4.9 n
FINM E OH U M t8-5.7 H
FINM REO MIW w7.7 n
0 '
b
Figure 3. O2 concentration in emulsion phase vs bed height at different temperatures.
11.5
I
i.5
BE0 tI$IGHT
Figure 5. O2 concentration profile in bubble phase.
-
m I
>
4c
.
0 N
0 N
Figure 6. O2concentration profile in bubble phase as a function of bubble diameter.
-
- - 1011
height, the carbon conversion, and the carbonate decomposition. Average Bubble Size. As shown in Figure 2 , the calculated bubble size varies widely depending on the theoretical equation used. Although the lack of experimental data makes difficult the assessment of these results, we point out the importance of grids in reducing bubble size. A small bubble size is also achieved with the operation of a small shallow multistage bed (Tamm, 1982). Oxygen Concentration. The variations in the oxygen concentration in the emulsion phase as a function of temperature, bubble size, and different bed depths are shown in Figures 3 and 4. As expected for a given temperature, the oxygen concentration is reduced as the bed height increases. This is because as the bed height increases, the bubble size increases. Hence, a smaller amount of oxygen is transferred to the emulsion phase. As the temperature increases, the rate of oxygen consumption goes up. Hence, for a given bed height, the lower the temperature, the higher the oxygen in the emulsion phase. Finally, for a given temperature (1200 O F ) , Figure 4 indicates that the oxygen concentration in the emulsion phase increases as the bubble size is reduced. The oxygen consumed by carbon, hydrogen, and sulfur in the emulsion phase is transferred from the bubbles, the size of which is a function of bed height, H (Table 11). Data presented in Figure 5 show that as the final bed height increases, the oxygen concentration in the bubble phase at a given distance from the air distributor increases. This is explained by the increase of bubble size and the reduced transfer of oxygen to the emulsion phase. This is confirmed by the calculated results shown in Figure 6, where the oxygen concentration in the bubble phase is shown to
___ 13 0 A I R f LB COKE
.
-- ._
15 0 LB A I R /LE CME
105-
3 4
I
RESIOENFE
6
I
I
I
TIME,~~INUTES"
I
14
16
Figure 7. Carbon burning efficiency at 1200 O F .
n
>. 100-
V w z
95-
E :
w
:
I
//
90-
I3
* z z a
85-
3
m
z
80-
0
m U
75-
RESIDENCE
TIRE, MINUTIES
12
I;
Figure 8. Carbon burning efficiency a t 1275 O F .
- 16 r.-
i
322 Ind. Eng. Chem. Res., Vol. 27, No. 2, 1988 Table VI. Decomposition Pressures for COz from MgC03, CaCO, atm temp, K MgC03 CaC03 894 29.27 0.004 922 82.19 0.0089 950 217.25 0.0174 978 543.43 0.0326
0
m A T L R E 74150 OF
-
TDFEUTURE T-1204
0
0 z
t , y)
40-
3-
zz U
g 201 a
Table VII. Variation in Carbon Conversion" pressure, C temD. K atm bubble model convers.. 7% 894 1.4 CGb 37 894 1.4 CG-5 grids 50 894 2.7 CG 62 894 2.7 CG-5 grids 100 978 1.4 CG-5 grids 64 978 2.7 CG-5 grids 100 a Average solids residence time = 10 min. bCG = Cranfield and Geldart (1972).
u U
=
I
1.1m
OF--EWIL. RSTAICTIM
7.1300
OF-WIL.rim"
E I
RESIDEd!
I
TIME $IN1
I
20
I
a
Figure 9. Carbonate decomposition.
-
3
t
decrease as the bubble diameter decreases. Effect of Excess Air on Carbon Combustion Behavior. The effect of excess air on the combustion is shown in Figures 7 and 8. Assuming that the retort temperature remains constant (798 K), these results demonstrate that the lower the combustor temperature, the lower the average solids residence time in the combustor zone. This coupled with the lower intrinsic reaction rate of carbon a t lower temperature (Table 111, a) results in lower carbon burning efficiency. However, the results in Figure 7 clearly show that adding excess air can raise the carbon burning efficiency to 100% provided that the solids residence time is greater than 8 min. Carbon Conversion: Other Process Variables. The variation in carbon conversion (percent carbon burned) as a function of combustor temperature, pressure, and bubble size is reported in Table VII. It should be restated that this model does not account for change of reactive particle diameter with time. As a result the calculated carbon conversions are conservative. It is, however, believed that the results shown in Table VI1 do indicate that a higher pressure operation of the combustor does result in increased carbon conversion and that the effect of bubble size depends on the unit pressure. At low pressures where the carbon burning rate is low, reducing the bubble size has relatively minor effect on carbon conversion. That is, the carbon conversion is kinetically controlled. A t high pressures where the kinetic rates are high, the effect of bubble size is significant. In this regime, the carbon conversion is mass transfer controlled. Carbonate Decomposition. The percent carbonate decomposition and the corresponding percent of heat of carbon combustion absorbed were calculated for the following case: Xc02,21 = 0.0609; Xco2,3,= 0.1048; pressure = 2.7 atm; combustor temperature = 894,922,950, 978 K; average solids residence time = 6-25 min. The percent carbonate decomposition (wed) is plotted as a function of residence time in Figure 9. The corresponding heat of carbonate decomposition (Whcd)as percent of heat of carbon combustion is plotted in Figure 10. The increase in Wcd and whcaat low temperatures with increased residence time is due to the increased rate of carbonate decomposition according to (1) and (2). At temperatures under 978 K, the decomposition of CaC0, is negligible as indicated by the corresponding rate constant (eq 2). This implies that a t low temperature, only MgCO, decomposes. This, as shown in Figure 10, has the advantage of low energy losses due to carbonate decom-
10-
0
.
WWATlAE 14240 OF
u 0
RSTAICTIM
1-1200 o F f W X L . RsTRICTID(
" E14250 OF
0
x
T-1150 oFfWIL.
OF
20
0
a
a m u t U W I
I
%Figure 10. Heat of carbonate decomposition
position. The flattening of the curves in Figures 9 and 10 a t high temperatures is due to the fact that all available MgC03 has been decomposed and the CaCO, contribution is very low a t these conditions. When the recombination of C02with MgO and CaO is taken in account, the percent carbonate decomposition is reduced. As shown in Figure 9, the reduction is greater at low temperatures and low residence times. in order to explain these results, the decomposition pressures of C02, shown in Table VI, should first be reviewed. At low temperatures where only MgC03 is decomposed, is much lower than at the decomposition pressure (Pcoz,es) high temperatures. Hence, at low temperatures, the rate of decomposition of MgCO, as calculated by (3) is lower than the rate of decomposition calculated from (1). At high temperatures where the decomposition pressure (Pco e ) is very high, the reduction in MgCO, decomposition e:$ to the presence of C02is negligible. These results imply that minimum carbonate decomposition (less than 15%) is achieved when the combustor operates at temperatures below 922 K and for residence times of less than 15 min.
Conclusions and Recommendations On the basis of this work, the following conclusions are drawn. 1. For the retorted, size distribution selected particles (average particle diameter 1600 pm), complete carbon conversion is feasible at high pressures (2.7 atm) and over the entire temperature range studied (894-978 K). 2. Bubble size was found to have an important effect, especially at conditions where reaction rates are high (high temperature and pressure). 3. The carbonate decomposition increases with combustor temperatures and residence time. Complete carbon conversion is feasible a t high pressures (2.7 atm) with less than 20% carbonate decomposition.
Ind. Eng. Chem. Res., Vol. 27, No. 2, 1988 323 4. At the preferred combustor operating conditions (high pressure, low temperature), the main reaction is dolomite decomposition, while calcite decomposition is negligible for the conditions considered in this study. 5. Recombination of COz with MgO occurs at low temperatures, high pressure, and long particle residence times. Based on the results of this study, there is a clear incentive to verify the conclusions of this study in a largescale system where pressurized combustion can also be tested. Realistic particle attrition data can then be generated, which can be incorporated into this model or an even more improved model.
Nomenclature A = bed cross-sectional area CcOzb= C02 concentration in bubble phase CcOze= C02 concentration in emulsion phase Cd = air distributor orifice coefficient COZb= O2 concentration in bubble phase COze= O2 concentration in emulsion phase COz,.= O2 concentration in combustion air C ,= molar heat capacity, air Cff = molar heat capacity, flue gas = heat capacity, spent shale D = fluid bed diameter Db = gas bubble diameter at height Z Db = mean bubble diameter in fluid bed Db,CG = mean bubble diameter according to Geldart and Cranfield (1972) model ijb,CGb =. mean bubble diameter according to Geldart and Cranfield (1972) model with five internal screens Dbm = maximum bubble diameter (Mori and Wen (1975) model) Dbo = initial bubble diameter (Mori and Wen (1975) model) Db,WM= mean bubble size (Mori and Wen (1975) model) Den, = oxygen diffusion coefficient in the spent shale particles do, = diameter of distributor openings d, = particle diameter d = particle diameter for class i = mean particle size for classes i and i + 1 E = activation energy Fa = combustion air rate F, = total carbon rate entering combustor Fco2,1= total amount of C02 produced from carbonate decomposition Fco = amount of C02 produced from decomposition of IdFCO, with one recycle only Fco ,31 = amount of C02 produced from decomposition of (?aC03 with one recycle only F f = flue gas rate FH = total amount of hydrogen in carbon residue Fa = total amount of pyrite sulfur in oil shale F, = rate of retorted solids entering combustor FR = total carbon combustion rate FR,I = rate of carbon burning in the retorted solids Fsh = oil shale flow rate to the retort F1 = recycle rate of spent shale to retort F2 = entrainment and withdrawn rate of particles out of the retort fi0 = mass of carbon burned per unit mass of oxygen consumed G = superficial air mass flux in the combustor g = gravitational constant H = bed height H,= heat of reaction: i = H, hydrogen; i = S, sulfur in pyrite; i = 2, MgC03;i = 3 CaCO,; i = 8, carbon combustion to COZ K* = Entrainment rate per unit combustor cross section K d = gas-solid mass-transfer coefficient K 2 = rate coefficient for MgC03 decomposition K 3 = rate coefficient for CaC03 decomposition KO= Arrhenius preexponential factor K = mass-exchangecoefficient between bubbles and emulsion $ = rate coefficient for c + o2reaction N = recycle number
4,
dpi
No,= number of distributor holes N,= dimensionless number, K&/Ub P = center-to-center distance between distributor holes People = C02 partial pressure in emulsion phase Pco,,~,= partial pressure of C02 produced from MgC03 P c o , , ~=~partial pressure of COPproduced from CaC03 Qcmb= energy absorbed by carbonate decomposition Q, = energy produced from carbon combustion QH = energy produced from H combustion Qs = energy produced from S combustion rl = mean particle radius for particle i i;, = mean particle radius for fluid bed R = gas constant Red = Reynolds number at minimum fluidization conditions Re, = Reynolds number at terminal particle velocity conditions R2 = rate of reaction of MgC03 decomposition R3 = rate of reaction of CaC03 decomposition R E = overall rate of reaction of C C02 RBavret = intrinsic carbon burning rate for retorting solids REa,,, = intrinsic carbon burning rate for spent solids Rsbl = diffusion rate of o2in the pores of particles of class i REcl= bulk diffusion rate of O2 from gas phase to particle surface REl,ret= rate of oxygen consumption for the retorted solids of class i Rsrap= rate of oxygen consumption for the spent solid particles of class i Rapt = total oxygen consumption rate for the retorted particles RE,,, = total oxygen consumption rate for the spent shale particles Sc = Schmidt number SFC02,21 = total C02production rate from MgC03 including recycle SFC02,31 = total C02 production rate from CaC03 including recycle T = dense bed temperature T, = combustion air temperature TB = base temperature for kinetic constants Tco,,~= total amount of C02 in solids entering combustor Tco2,21 = total amount of C 0 2 in MgC03 in solids entering combustor Tco,3i = total amount of C02 in CaC03 in solids entering combustor TpR = oil shale preheat temperature TR = retorting temperature ub = bubble rise velocity umf= minimum fluidization velocity umf,= minimum fluidization velocity for particles of class i u,, = velocity of gas through an orifice in distributor ut = particle terminal velocity u, = gas superficial velocity u: = initial concentration of kerogen in shale U2 = concentration of COPas MgC03 in retorted shale particles U3= concentrationof C02as CaC03in retorted shale particles @ = carbon concentration in retorted solids U, = carbon concentration in spent solids W,d = percent of carbonates decomposed whd = percent of heat of combustion consumed by carbonates decomposition Wret = weight fraction of retorted solids in fluid bed W,, = weight fraction of spent solids in fluid bed X , = weight fraction of carbon in spent solids X , = weight fraction of carbon in retorted solids X c o 21 = weight fraction of C02 as MgCO, in the retorted sojids
-
I n d . E n g . Chem. Res. 1988,27, 324-328
324
Xco2,3i= weight fraction of C02as CaCO, in the retorted solids Xco2,20= weight fraction of C02 as MgCO, in the spent solids Xco2,30= weight fraction of COPas CaC03in the spent solids XH = weight fraction of H in carbon residue
Xi= weight fraction of particles in class i
X s = weight fraction of S in retorted solids 2 = distance from air distributor Greek Symbols h p b = bed pressure drop h p d = distributor pressure t = void fraction t b = bubble fraction
drop
emf = void fraction at minimum fluidization conditions p = gas velocity pg = density of gas ps = particle density
4 = particle sphericity
Literature Cited Baughman, G. L. Synthetic Fuels Data Handbook; Cameron Engineers: New York, 1978; p 14. Braun, R. L. Report UCRL-53119; 1981; Lawrence Livermore Laboratory, London. Campbell, J. H. 1978 Report UCRL-52089, Part 11, 1978; Lawrence Livermore National Laboratory; London.
Cranfield, R. R.; Geldart, D. Chem. Eng. J . 1972, 3, 211-231. Fluidization Engineering; Robert E. Krieger: Huntington, NY, 1977. Geldart, D. EPRI CS-2094, Oct 1981; EPRI, Washington, DC. Hall, R. N., AIChE Meeting, Anaheim, CA, June 1982. International Critical Tables; McGraw-Hill: New York, 1930; Vol. VII, p 292. Knowlton, T. M., AIChE Meeting, New York, Dec 1-5, 1974. Mori, S.; Wen, C. Y. AIChE J. 1975, 2I(1), 109-115. Patzer, J. F.; Carr, N. L., Presented at the AIChE National Meeting, Anaheim, CA, 1984. Rammler, R. W., Oil Shale Processing Technology; Dean Allred, V., ed.; Ellenar Graphics: New York, 1982. Sohn, H. Y.; Kim, S. K. Znd. Eng. Chem. Process Des. Deu. 1980,19, 550. T a " , P. W. "Combustion of Pyrolyzed Carbon Containing Solids in Staged Turbulent Bed". U S . Patent 4 336 128, 1982. Ta", P. W.; Bertelsen C. A.; Handel, G. M.; Spars, B. G.; Wallman, P. H. Energy Prog. 1982,2(1), 37. Vasalos, I. A.; Tatterson, D. F.; Furlong, M. W.; Kowalski, T. L.; So, B. Y. C., Annual AIChE Meeting, San Francisco, 1984. Wallman, P. H.; Tamm, P. W.; Spars, B. G. Prepr.-Am. Chem. SOC., Diu. Pet. Chem. 1980, 25(3), 70. Wen, C. Y.; Yu, Y. H. AIChE J . 1966,12, 610.
Received for review February 21, 1986 Revised manuscript received May 27, 1987 Accepted October 8, 1987
SEPARATIONS Absorption of Carbon Monoxide into Aqueous Solutions of K2C03, Methyldiethanolamine, and Diethylethanolamine C. J. Kim,* Alan M. Palmer, and George E. Milliman E x x o n Research and Engineering C o m p a n y , Corporate Research Laboratories, Annadale, N e w Jersey 08801
Rates of CO absorption into 1 M aqueous solutions of K2C03,diethylethanolamine (DEAE), and methyldiethanolamine (MDEA) were measured in a stirred tank reactor a t 348-398 K and a t CO pressures in the range of 7.5-31 bar. The absorption behavior showed the absence of any mass-transfer limitations, and detailed kinetic analyses established that a reaction step, CO + OHHCOO-, is rigorously controlling the absorption rates in all systems examined in this study. Comparative discussions on the kinetics of C 0 2 OH- vs CO OH- are given.
-
+
Removal of C02and H2Sfrom gases is usually achieved by regenerative absorption into -aqueous solutions containing bases such as amine-promotedpotassium carbonate solutions and aqueous amines. In cases when the feed gas contains carbon monoxide, the gas-treating solutions become gradually deactivated due to the formation of stable formate salts. A brief account of such deactivation of diethanolamine-promoted potassium carbonate solutions in commercial plants was described by Eickmeyer (1962). The reaction of CO with K2C03 has been shown by Yoneda et al. (1943) to occur by CO + K2C03+ H 2 0 KOOCH + KHC03 (1) According to the thermodynamic data compiled by Latimer (1952), the standard free energy of reaction 1is -19.2 kJ/mol, while that of reaction 2 is -14.4 kJ/mol. Under
-
*Present address: 1240 Rattlesnake Bridge Road, Bedminster, N J 07921. 0888-5885/88/2627-0324$01.50/0
+
COZ
+ K&03 + H2O
+
2KHCO3
(2)
the usual gas-treating conditions, regenerative CO desorption by the reverse process of reaction l cannot be expected because the facile COz desorption via the reverse step of reaction 2 depletes HC03-, the proton-donating species needed for decomposition of the formate ion. It is thus expected that CO absorption leads t o degradation of the basic components of the gas-treating solutions into nonregenerable formate salts. The extent and severity of such degradation depend on the rates of CO absorption, and the present study was conducted to define the basic chemistry involved in the reactions of CO with K,CO,, methyldiethanolamine (MDEA), and diethylethanolamine (DEAE) in aqueous solutions. I
I.
Experimental Section Absorption Rate Measurement Procedure. A 300cm3 stirred autoclave (Autoclave Engineers Hastloy-C 0 1988 American Chemical Society
