UCRL-53517

Process Simulation Model for a Staged, Fluidized-Bed

Oil-Shale Retort With Lift-Pipe Combustor

J. C. Diaz

R. L. Braun

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LPT ..&

February 1984

* . - •

CIRCULATION COPi SOWECT TO RECALL

IN TWO WEEKS

DISCLAIMER

This document was prepared as an account of work sponsored by an agency of the United States Government. Neither the United States Government nor the University of California nor any of their employees, makes any warranty, express or implied, or assumes any legal liability or responsibility for the accuracy, com­pleteness, or usefulness of any information, apparatus, product, or process disclosed, or represents that its use would not infringe privately owned rights. Reference herein to any specific commercial products, process, or service by trade name, trademark, manufacturer, or otherwise, does not necessarily constitute or imply its endorsement, recommendation, or favoring by the United States Government or the University of California. The views and opinions of authors expressed herein do not necessarily state or reflect those of the United States Government thereof, and shall not be used for advertising or product endorsement purposes.

Work performed under the auspices of the U.S. Department of Energy by Lawrence Livermore National Laboratory under Contract W-7405-Eng-48.

UCRL-53517 Distribution Category UC-11

Process Simulation Model for a Staged, Fluidized-Bed

Oil-Shale Retort With Lift-Pipe Combustor

J. C. Diaz

R. L. Braun

Manuscript date: February 1984

LAWRENCE LIVERMORE NATIONAL LABORATORY IP University of California • Livermore, California • 94550 | g T

Available from: National Technical Information Service • U.S. Department of Commerce 5285 Port Royal Road • Springfield, VA 22161 • $7.00 per copy • (Microfiche $4.50 )

Contents

Nomenclature iv Abstract 1 Introduction 1 System Simulated 2

Overview 2 Retort 3 Combustor 3 Overall System 3

Modeling of System 3 Retort 3

Chemical Processes Considered 4 Physical Processes Considered 6 Mass and Energy Balance 6

Combustor 7 Chemical Processes Considered 8 Physical Processes Considered 10 Mass and Energy Balance 12

Auxiliary Equipment 14 Steady-State Solution of System 15

Application of the Model 15 Case Simulation 15

Acknowledgments 17 References 18 Appendix 19

in

Nomenclature

C0 = Oxygen concentration (kg 0 2 / m 3 gas)

Cp co = Specific heat of C 0 2 in surge bin 0/kg-K) C = Specific heat for the gas phase in the combustor 0/kg • K) C ; = Specific heat for fth species in the gas phase (J/kg • K) Cp: = Specific heat for the ;'th particle in the combustor Q/kg • K) C K = Specific heat for organic material in shale (J/kg • K) C N = Specific heat for inorganic material in shale (J/kg • K) C 0 j = Specific heat for 0 2 , between the gas and solid temperatures (J/kg- K) dj = fth particle diameter (m) Dc = Combustor diameter (m) D = Harmonic mean particle diameter (m) fov = Fraction of oil in vapor phase, from kerogen pyrolysis /i = Stoichiometry factors for kerogen pyrolysis (kg/kg kerogen) FB = Bitumen concentration ( k g / m 3 shale) FK = Kerogen concentrat ion ( k g / m 3 shale) FK = Original kerogen concentrat ion ( k g / m 3 shale) FN = Concentrat ion of inorganic material in r aw shale ( k g / m 3 shale) Fw = Concentrat ion of b o u n d water in raw shale (kg H 2 0 /m 3 shale) ga = Gravitational acceleration constant (m/s ) G; = Gas phase mass flux of fth species (kg/m2 • s) GCQ2 = Gas phase mass flux of C 0 2 (kg/m2-s)

GH 0 = Gas phase mass flux of H 2 0 (kg/m2-s)

GN = Gas phase mass flux of N2 (kg/m2-s)

G0 = Gas phase mass flux of 02 (kg/m2-s)

Gs = Solids mass flux (kg/m2-s) h: = Gas-solid heat transfer coefficient (J/m • K • s) I = Maximum allowed change of carbon concentration in

particle ;= 1, in combustor (kg C/m shale) N = Number of stages in retort NRe = Reynolds number NSc = Schmidt number NSh = Sherwood number Q: = Total heat of reaction of ;'th particle in combustor Q/m3 bed-s) r- = /th particle radius in combustor (m) RB = Rate of change of bitumen concentration (kg/m3 shale • s) R = Overall carbon oxidation rate (kg C/m3 shale-s) RD = Rate of dolomite decomposition (kg C0 2 /m 3 shale • s) R H . = Overall hydrogen oxidation rate (kg H/m3 sha les ) RK' = Kerogen reaction rate (kg/m3 shale-s) RMj = Rate of calcite decomposition (kg COz /m3 shale-s) RQ . = Overall oxygen reaction rate, /th particle (kg 0 2 /m 3 shale • s)

Rn = Intrinsic chemical reaction rate (kg 0 2 / m 3 shale • s) u 2 C

Ro2D = Intraparticle 02 diffusion rate (kg 02 /m* shale-s)

R0 = Gas-solid mass transfer rate (kg 0 2 /m 3 sha les )

Rs. = Overall sulfur oxidation rate (kg S/m3 sha les ) Rw = Rate of release of bound water in raw shale (kg/m3 shale • s) Rz = Rate of char production (kg/m3 sha les ) S = Raw shale processing rate in retort (kg/m • s) tr = Solids residence time in retort (s)

iv

T = Gas phase temperature in combustor (K) T: = Temperature of /th particle (K) Tr = Reference temperature for enthalpy calculations (298 K) U„ = Gas linear velocity (m/s) L7j = Local linear velocity for /'th particle (m/s) L7S j = Local slip velocity of /'th particle (m/s) Ut j = Local free terminal velocity of /'th particle (m/s) V-. = Volumetric flow rate of /'th particle (m3/s) Wc . = Concentration of organic carbon in /'th particle (kg/m3 shale) WD = Concentrat ion of dolomite C 0 2 in / th particle, minus calcite

C 0 2 generated from dolomite decomposi t ion (kg C O z / m shale) V V H : = Concentration of hydrogen in /th particle (kg/m3 shale) WM: = Concentration of calcite C 0 2 in /'th particle, plus calcite

C 0 2 generated from dolomite decomposition (kg C0 2 /m shale) W0 = Stoichiometric oxygen required to oxidize remaining reactive

2 material in shale, in combustor (kg 0 2 / m shale) W0 0 = Stoichiometric oxygen required to oxidize original reactive

material in shale, in combustor (kg 0 2 / m shale) Ws. = Concentration of sulfur in /'th particle (kg/m3 shale) / = Gas-wall Fanning friction factor / = Solid-wall friction factor AHC = Heat of reaction for carbon combustion (J/kg Q AHD = Heat of reaction for dolomite decomposition (J/kg C02) AHH = Heat of reaction for hydrogen oxidation (J/kg H2) AHK = Heat of reaction for kerogen pyrolysis (J/kg kerogen) AHM = Heat of reaction for calcite decomposition (J/kg C02) AHS1 = Heat of reaction for sulfur oxidation at T > 650°C (J/kg S) AHS2 = Heat of reaction for sulfur oxidation at T < 650°C (J/kg S) AHW = Heat of reaction for bound water release (J/kg water) Az = Height of differential element in combustor (m) «j = Packing fraction of /th particle (m3/m3 bed) ar = Volume ratio of recycled solids to raw shale (m3/m3 raw shale) « r = Bed void fraction (m3 v o i d / m 3 bed) pg = Gas phase density ( k g / m 3 gas) p- = Density of /'th particle in combustor ( k g / m 3 shale) Pr = Raw shale density (kg /m 3 shale) ps = Density of recycle solids stream (kg /m 3 shale)

v

Process Simulation Model for a Staged, Fluidized-Bed

Oil-Shale Retort With Lift-Pipe Combustor

Abstract

We have developed a computer model to simulate an aboveground oil-shale retorting process that utilizes two reactors (a staged, fluidized-bed retort and a lift-pipe combustor). This model calculates the steady-state operating conditions for the retorting system, taking into account the chemical and physical processes occurring in the two reactors and auxil­iary equipment. The release of mineral water and the pyrolysis of kerogen take place in the retort when raw shale is mixed with hot partially-burned shale, and the partial com­bustion of residual char and sulfur takes place in the combustor as the shale particles are transported pneumatically by preheated air. The retort is modeled as a series of stirred-batch reactors, and the combustor is simulated using a lumped-parameter model of finite-difference elements. Simulation results include stream flow rates, temperatures and pres­sures, bed dimensions, and heater, cooling, and compressor power requirements.

We have already used the model to simulate a hypothetical commercial operation to produce an oil-plus-gas equivalent of 50,000 bbl/day, using 30-gal/ton shale. Although data for model validation are not yet available, the calculated results of the simulation appear reasonable.

Introduction

Lawrence Livermore National Laboratory (LLNL) is involved in the development of oil-shale retort­ing technology. Ongoing work includes an experimental effort that focuses on the chemistry of oil shale,1,2

the operati on of pilot-scale retorts,3 and the development of computer models to simulate oil shale retort­ing.4 The latter is of particular value because it incorporates the most current information available; hence it assesses our understanding of retorting processes and helps reveal areas where additional research is necessary. The LLNL one-dimensional computer model, originally developed for modified in-situ retort­ing,3 has been modified continually to handle aboveground retorting processes that use a moving, packed bed. They include hot-gas, internal-combustion, and cascading-bed retorts.

The use of stirred-bed retorts along with fluid-bed combustors (e.g., Chevron5 and Lurgi6) has been considered for several processes, but no known computer models for simulating these processes have been made available to the public. Therefore, a model (separate from the packed-bed model), to simulate a process that uses a staged, fluidized-bed retort coupled to a lift-pipe combustor, has been under devel­opment. This computer model encompasses a number of chemical and physical processes that are ana­lyzed in terms of the appropriate kinetics and thermodynamics involved. The development of the model has already proved valuable in defining areas where much research is necessary.

1

System Simulated

Overview

We have developed a computer model to simulate a retorting process whereby raw shale is mixed with hot burned shale to induce kerogen pyrolysis. The retorted shale is then combusted in the presence of air to provide the hot shale. The process we attempt to simulate uses two reactors in a loop configura­tion, as shown in Fig. 1. The first reactor is a fluidized bed whose principal function is to allow the raw and burned shale feed streams to mix thoroughly. It is also required to provide the residence time necessary for kerogen pyrolysis. To provide the fluidizing gas for this reactor, a noncondensable fraction of the off gas is recycled. When the mixed stream of solids leaves the retort and enters the second reactor (lift-pipe combustor), combustion takes place as the solids are lifted pneumatically. In addition to calculating stream conditions and flow rates, the model also calculates approximate reactor dimensions.

Equipment peripheral to the reactor loop must also be considered when estimating the stream condi­tions and energy requirements of this retorting system. Such equipment includes a condensing system for the gas and vapor products from the retort, as well as heaters and compressors for the gases needed at each reactor. These units are modeled simply because, at this point, no valuable information would be gained from a more detailed simulation. The retort and combustor, in contrast, have been modeled rigorously so that we may identify and understand the more important process parameters.

Raw shale Flue gas

Gas products **

Foul water • *

Oil products

Figure 1. Schematic of fluidized-bed retort system.

2

Retort

Many models are available for simulating fluidized-bed reactors, and a number of them could have been modified for our purpose. The modifications would have been extensive though because those models ir corporate an entirely different chemistry and account for items (e.g., internals and economics), that do mat apply here.

The goal in developing a retort model is to include those physical and chemical processes that have the greatest effect on the kinetics and energy distribution within the retort. Toward this goal, we envision that the solid material entering the retort quickly mixes and travels downward in the retort through a series of stages that reduce vertical mixing in the retort and induce a plug-flow effect on the solids stream. The fluidizing gas travels in a similar plug-flow manner, countercurrent to the solids stream. Within each stage, the model accounts for the presence of solids, kerogen, oil vapor, char, water vapor, CHX, CH4, H2, C0 2 , and CO. Bound water is considered only in the first stage. The model also accounts for the possibility that raw shale fines will be elutriated from the retort. In such a case, we assume that those fines undergo pyrolysis within the gas stream at the retort exit and that the streams reach thermal equilibrium before being separated at the cyclones. The fines and solids stream leaving the retort are then fed to the combus-tor. The retort can be modeled with 2 to 20 stages, using any one of various kerogen pyrolysis models.

Combustor

In the combustor, the particles of just-retorted and partially-burned shale are entrained by preheated inlet air and lifted pneumatically at velocities calculated as a function of particle properties and process conditions. Each particle undergoes combustion and carbonate decomposition at a rate prescribed by the local chemical and thermodynamic conditions, taking into account mass- and heat-transfer limitations. The reactive species include undecomposed kerogen, residual organic carbon and hydrogen, pyrite, pyr-rhotite, dolomite, and calcite. The reaction products are carbon dioxide and water vapor. We assume that the combustion of any undecomposed kerogen is not hindered by mass-transfer limitations.

At the combustor exit, the solids are separated from the flue gas and reach an equilibrium tempera­ture equal to the mass-averaged temperature of the individual particles. After the fines have been sepa­rated, the desired quantity of hot shale is recycled into the retort.

Overall System

The results of the reactor and equipment calculations include flow rates, temperatures, and pressures for the gas streams and flow rates and temperatures for the solids streams. Compressor power require­ments and heater and cooling loads are also calculated for peripheral equipment. These results are inte­grated by the model in a manner consistent with the flowsheet, and an overall mass and energy balance is calculated for the system. For the mass balance, system input streams include raw shale fed to the retort and air used for combustion. System output includes burned shale, flue gas from the combustor, shale oil, noncondensable gases, and condensed water. The energy balance includes the sensible heats of the input and output streams, heats of reactions, heater and cooling loads, and compressor power requirements.

Model ing of System

Retort

Retort modeling is carried out in a simple, stepwise manner. The chemical and physical processes considered, in modeling the retort are listed in Table 1, together with information input to the model. The retort itseli: is modeled as a series of adiabatic countercurrent stirred-batch reactors into which two particle

3

Table 1. Processes considered and input to the retort model.

Chemical and physical processes Kerogen pyrolysis Bound water release Fhiidization of solids mixture Bed pressure drop

Input to retort model Raw shale inlet temperature Raw shale grade and composition Fluidizing gas temperature Pyrolysis reaction rate order Residence time or degree of pyrolysis Processing rate Burned shale recycle ratio Particle size distribution Number of stages to simulate

size distributions are input—one to describe raw shale, the other recycled shale. An additional particle description can be used if fines exist in the system. A mass-balance solution is made simple by assuming that there are no reversible chemical reactions; therefore, given the residence time and an estimated temperature, a mass balance is evaluated for all species within each stage. If the residence time is not given, the model estimates the time and calculates the degree of kerogen pyrolysis that would result, compares the calculated degree of pyrolysis with that given in the input, and adjusts the estimated residence time. This process is continued until the desired degree of pyrolysis is attained. Then the mass balance is evaluated.

Solving the energy balance for the retort is not so simple because of the split-boundary conditions given. The problem is solved by writing a linear, first-order equation for the energy balance at each stage, then solving all the equations simultaneously via the Gauss-Jordan reduction method. Interparticle heat transfer has not been included in the retort simulation at this point, so the gas and solids streams are assumed to reach thermal equilibrium at each stage. The model calculates the temperature and flow rates of the gas and solids at the inlet and outlet of each stage. It also estimates the pressure drop and bed dimensions for the retort.

Chemical Processes Considered Kerogen pyrolysis can be simulated with one of two models. If the model presented by Wallman

et al.7 is chosen, the kerogen pyrolysis rate may be expressed as

RK = )t1FK . (kg/m3 shale-s)

The rate at which the bitumen concentration changes is

and the rate of char production is

^Z = /c*3^B '

where

it, = 9.63 X 1010 exp(-21940/T) , (s_1)

k2 = 3.0 X 103 exp(-11370/T) , (s_1)

*3 = 3.98 X 103 exp(670dj) exp(-11370/T) , (s_1)

4

and the stoichiometric factors

/b = 0.2612 (kg bitumen/kg kerogen)

and

fc = 0.80 . (kg char/kg bitumen)

The rate constants k2 and k3 are, respectively, the heavy oil and coke production constants. Kerogen pyrolysis can also be expressed as a single Mth-order reaction,

RK = k FK0 [j^ J • (kg/m3 shale • s)

Fluidized-bed experiments8 were consistently interpreted9 with n = 1.4 and k = 6.9 X 1010

exp( —21790/T) {s~l). Conventional nonisothermal pyrolysis experiments,10 on the other hand, were con­sistent with w = 1 and k = 2.81 X 1013 exp(-26390/T) (s - 1) . The endothermic heat of reaction, recom­mended by J. F. Carley11 for kerogen pyrolysis, is

AHK = 3.30 X 105 . (J/kg kerogen)

The reaction rate for the release of bound water is assumed to be fast enough to allow complete release within the first stage of the retort. Thus,

Rw = kwF w , (kg H 2 0 /m 3 shale • s)

where

kw = instantaneous at T > 400°C (s_1)

and

AHW = 2.26 X 106 . (J/kg H20)

The subsequent production of char, oil, and gases from kerogen pyrolysis for an «th order reaction is calculated by means of stoichiometric factors given by Burnham et al.12 They are:

fx = 0.2025 (kg residual char/kg kerogen)

f2 = 0.6767 (kg oil/kg kerogen)

f3 = 0.0359 (kg C0 2 /kg kerogen)

/4 = 0.0057 (kg CO/kg kerogen)

/5 = 0.0268 (kg H 2 0/kg kerogen)

f6 = 0.0030 (kg H2/kg kerogen)

f7 = 0.0142 (kg CH4/kg kerogen)

/g = 0.0352 . (kg CHx/kg kerogen)

5

Then, for this case, the rate of char production is

Rz = R-Kn (kg char/m3 shale • s)

and the rate at which the bitumen concentration changes is

RB = 0.0 .

Physical Processes Considered The superficial gas velocity in the retort is calculated with a modified version of the Ergun equation.

This equation was modified by D. Christiansen13 on the basis of experimental results for the fluidization of crushed, raw shale.13 The modified equation is of the form

A-N2Re + B.NRe + C =0.0 ,

where

3.56 A =

63

g _ 620(1 -<) t3

r -Pj>gga(Ps-*>g)

and

e = bed void fraction = 0.60 (m3 void/m3 bed) .

The pressure drop across the bed of solids is estimated from the pressure produced by the force of gravity on the solids within the bed. The actual pressure drop is assumed to be 1.8 times greater than that estimatecl because of the effects of the main gas distributor and redistributors that would exist within a staged, fluidized bed.

Mass and Energy Balance The mass-balance equations for each stage can be written in finite terms, using the reaction rate

expressions and stoichiometric factors shown above. A mass balance for the solids may be written at the mth stage, with stages S. < m, where solids flow

from $ to m. The balance in terms of (kg/m2 • s) is

(jj (fK,m + FN,m + Ps«r) = (jj ( F w + FN>, + PsaT - RKmAt + Rz,mAt) ,

where

At =••— . (s) N

6

A mass balance for each species of the gas stream may be written at the mth stage, with stages m < n, where gases flow from n to m. The balance in terms of (kg/m2 • s) for the ith species is

Gi,m = G i # n - | -RK ,mW^j .

Given the above equations, the overall mass-balance equation for the mth stage is

Because thermal equilibrium is assumed to exist between the gas and solids streams leaving a stage, the energy-balance equation at each stage encompasses both the gas and solids phases. The overall energy balance is based on equating the enthalpies of the streams, relative to 298 K, by using mean heat-capacity values.

The general equation for an adiabatic system is

AH(reaction) + AH(products) - AH(input) = 0.0 ,

where, for this case,

AH(reaction) = R K m AtAH K ( - j , (J/m2-s)

AH(products) = f-)[FK,mCpXm(:rm - Tr) + (FN m + psar)Cp:Nm(Tm - Tr)] + ^G i /mCp, i im(Tm - Tr) ,

and

AH(input) = | - j [ F M C p X j f (Tg - Tr) + (FNJ) + psat)CpjNi((Te - Tr)] + ^GlnCpXjTR - Tr) .

Combustor

Table 2 is a listing of the chemical and physical processes accounted for by the combustor model, together with information input to the model. The chemical reaction rates of oxidation of residual organic carbon, residual organic hydrogen, and mineral sulfur species under conditions of a lift-pipe combustor have not been determined as yet. Therefore, the same reaction rate is used for all three reactions. The chemical reaction rate measured by Sohn and Kim14 is used here, but the rate coefficient is increased by a factor of 10. This modification is based on preliminary observations by R. Taylor15 during experiments on the combustion of just-retorted shale.

The combustor simulation calculates physical properties for each particle described as a function of height in the combustor, but no axial dispersion or radial gradients are considered. Particle attrition is implied by accounting for two particle size distributions of three groups each, one for just-retorted shale and one I'or recycled shale. An additional particle description can be used to account for the presence of fines in trie system. The conditions given pose an initial value problem that allows the simulation to be solved as a lumped-parameter model of finite-difference elements, each of which is analyzed twice. In the first itera tion, all the thermodynamic properties and governing equations at each element are solved using the exit conditions of the previous element. In the second iteration, the calculated conditions at the present element are averaged with the exit conditions of the previous element, and the calculations are repeated.

7

Table 2. Processes considered and input to the combustor model.

Chemical and physical processes Combustion of residual organic carbon and hydrogen Combustion of pyrite and pyrrhotite Combustion of nonpyrolyzed kerogen Decomposition of dolomite and calcite Pneumatic transport Heat transfer between solids and gas streams Pressure drop and change in void fraction Particle attrition

Input to combustor model Air temperature Oxygen stoichiometric ratio Minimum void fraction at combustor bottom Initial air velocity incremental factor Incremental change in char concentration of largest particle Combustor exit pressure Solids temperature desired or combustor height Particle size distributions Surge bin residence time

These results are then used as input conditions at the next element. The element size A z is regulated by the incremental change in organic carbon concentration of the largest, just-retorted shale particle. Note that this incremental change in concentration is an input parameter. Because of the nonlinearity of the process, A z is not constant, but the relative number of elements analyzed may be controlled. Also note that model calculations do not simulate the acceleration region that exists in pneumatic transport systems; hence, the implication is that particles approach an established flow velocity upon entering the combustor. This weakness in the model tends to overstate the required combustor height.

In simulating the combustor, one has the option of using a surge bin to retain the solids leaving the combustor. As listed in Table 2, the only parameter for bin simulation is the residence time. The bin is considered a packed, plug-flow reactor, with carbonate decomposition being the only reaction taking place. If the specified residence time is greater than zero, the bin is simulated as a series of 50 equivolume batch reactors.

Chemical Processes Considered The combustion rates of char and iron sulfide take into account gas-solid mass transfer, intraparticle

diffusion, and intrinsic chemical reaction rates for each particle described. The gas-solid mass transfer rate is

Ro2M = —~ '

')

the rate of intraparticle 02 diffusion is

- l

and the intrinsic chemical reaction rate is

R _ t P ^ Wj, 2M

8

Re 3DeO,Q) 'Wo2oY / 3

-'2D

(kg 02/m3 shale • s)

where

Wo, = VJ2

wc WH ws -7^ + -r- + -r1

V/c /H / S . (kg 02/m3 shale)

and the oxygen stoichiometric factors are

/c = 12/32 (kg C/kg 02)

/H = 2/16 (kgH/kg02)

and

/s = 64/144 . (kgS/kg02)

The combined effect of the above rates is approximated from the law of additive reaction times16 as

°2'' 1 1 1 + ^ + Ro2C

Ro2D *C

This overall rate expression for the consumption of oxygen can be distributed proportionally for each reactive component in the shale. For the combustion of organic carbon in the jth particle,

RCi = Ro2,i(^1) • (kg C/m3 shale-s)

The reaction rates applied to hydrogen and sulfur (RH j and Rs j) are similarly defined. In the above equations, the chemical reaction rate coefficient is calculated locally as

kc = 24.7exp(-11100/T,) . (s^Pa^1)

The diffusion coefficient17 is expressed as

DeO2 = 1.04Xl0-15(FKo)2Tj165 , (m2/s)

and the mass transfer coefficient (kd) is estimated from the Sherwood number,

Nsh = 2 + 0.6NSc]/Xe1/2 •

The exothermic heats of reaction for the combustion of organic carbon and hydrogen are, respectively,

AHC = -3.28 X 107 a/kg C)

and

AHH = - 1 2 1 X 108 . (J/kgH2)

The chemistry of sulfur in shale is currently being investigated; thus, the model accounts for sulfur in a simple manner. Burnham and Taylor18 indicate that the original sulfur concentration for Anvil Points shale is approximately 0.60 wt%, but that this concentration decreases to 0.39 wt% (based on the mass of raw shale) after pyrolysis. They also state that the oxidation of sulfur may follow any number of reaction paths, depending on the local conditions and nature of the sulfur. In the model the sulfur content of the

9

just-retorted shale is assumed to be 0.39 wt%. To simplify the chemistry, only two possible reaction paths are assumed to exist, with the appropriate path being dictated by the particle temperature.

At a temperature below 650 °C, the important reaction is

2FeS + CaMg(C03)2 + 9/2 02 -, Fe203 + CaS04 + MgS04 + 2C0 2 ,

with an exothermic heat of reaction of

AHS1 = -2.802 X 107 . (J/kg S)

The second possible reaction path, for temperatures above 650°C, is

2FeS + 2CaC03 + 9/2 02 — Fe203 + 2CaS04 + 2C0 2 ,

with the exothermic heat of reaction being

AHS2 == -2.905 X 107 . G/kg S)

The model accounts for the decomposition of carbonate minerals in a simplified way.4 It considers two foi-ms of mineral C0 2 . One form is for carbonates such as dolomite that decompose at lower tempera­tures. The second accounts for the calcite originally in the shale and that generated from dolomite decom­position. The reaction rates for dolomite and calcite decomposition are, respectively,

RD. = A:DWDj (kg C0 2 /m 3 shale-s)

and

V j = M V j ' (kg C0 2 /m 3 shale-s)

where the rate coefficients are

kD = 1.7 X 1010 exp(-29090/T) (s_1)

and

Jtv, = 9.6 X 101 0exp(-36050/r) (s"1)

and the corresponding endothermic heats of reaction are

AHD = 3.0 X 106 (J/kg C02)

and

AHM = 2.9 X 106 . (J/kg C02)

Physical Processes Considered Most literature on the pneumatic transport of solids deals with uniform particle characteristics and

offers empirical solutions to evaluate mean process characteristics. However, for the given situation, we feel it more important to deal with the individual particles before evaluating their aggregate effect on the system.

To determine the initial air velocity at the combustor, we first calculate the amount of stoichiometric oxygen required to oxidize all organic carbon in the just-retorted shale. Then we increment or reduce this value by a factor input to the model. Once we know the mass and volume flow rate of air, we can choose a combustor diameter that will cause the air velocity to equal the greatest slip velocity of the particles involved. We then increase the air velocity by another factor input to the model and calculate the cor-

10

responding combustor diameter. At this point, we check the resulting void fraction against the minimum void fraction allowed (input to the model). If the criterion has been met, we calculate the particle veloci­ties. If, the ciiterion has not been met, we increase the gas velocity until the calculated void fraction is equal to or greater than the minimum allowed. In subsequent calculations of air velocity, we need account only for temperature, pressure, bed void fraction, and mass flow rate.

To calculate particle slip velocities, hence linear velocities, we had to apply a number of assumptions to the results; of experiments conducted by R. Quong19 on the vertical pneumatic transport of crushed shale. Quong determined particle slip velocities experimentally, as a function of particle size and mass flux, at ambient temperature and pressure. To apply these results to the model, we derived a correlation that evaluates the slip velocity from the mass flux, the particle diameter, and the particle terminal velocity calculated at local conditions. Model calculations then imply that the terminal-to-slip velocity relation­ships based on Quong's observations at ambient conditions remain constant even at the expected operat­ing conditions of a lift-pipe combustor and that the relationships are independent of wall effects.

The calculated terminal velocity is based on Newton's law,

^ 4rfj(Pj-Pg)ga

3pf;Q

where Cd is the value of the particle drag coefficient, as calculated from equations based on work by

Pettyjohn and Christiansen.20 The terminal-to-slip velocity correlation is

U, UU

s,J

100 + 1.41 tanh [(325 - 2.5GS) dj]

(m/s)

thus, the particle linear velocity is simply

U: U„ - U, s.J (m/s)

Pressure drop within the combustor is modeled in a manner recommended by Kunii and Levenspiel,21 which accounts for three sources of pressure drop: the static head, gas and solids-to-wall friction losses, and pressure loss caused by acceleration of the solids. The pressure-drop equation for the nth finite-difference element is

ybz) U / s + fe /g ^-A 5z

AM ,+2 (N/m3 bed)

where the static head is expressed as

,fe, = [ps(l-e) + Pge]ga ,

and the pressure drops caused by gas friction, solid-wall friction, and acceleration of the ;'th particle are, respectively,

D„

8P\ 52 /p,j

x/sLZgL7j ( 1 - 0

(*>«)

0.5

X'

11

and

( $ - ( < * . - < 4 - , *

In the eibove equations, the Fanning friction factor is evaluated from

/g = 0.0008 + 0.0552NRe-0237 ,

and the particle friction factor is evaluated from a correlation derived for a plot given by Kunii and Levenspiel.22 The equation is

/Sfj = 0.046 W R e i -a 5 6 9 .

Because of the changes in the volume flow rate of the gas and particle velocities, the void fraction within the combustor must be evaluated at each element. Evaluation is accomplished by

m e = 1 _ . (m3 void/m3 bed)

fxD2X

Mass and Energy Balance

The mass and energy balance at each combustor element are evaluated twice. As mentioned, in the second iteration for an element mean values are used for the pertinent parameters. On the basis of the rate expressions given above, a mass-balance equation may be written for each particle described.

Before the first iteration through an element, we calculate the height of the element from the esti­mated reaction rate for the largest just-retorted particle, the maximum allowed change in carbon con­centration (I), and the particle linear velocity (Ux). If the residence time for the largest particle is

Afj = I /R C 1 , (s)

the element height is

Az = UXMX (m)

and the residence time for each of the other particles is

Afj = Az/Uj . (s)

At any axial point along the combustor, the mass balance for the ;'th particle may be expressed as

(kg/m3-s) ^

Similarly, at any point along the combustor axis, we write a mass-balance equation for each gas specie (kg/nv3 • s). The equations are

(5GA I '-• I = 0.0 for nitrogen ,

12

rH - T. V iz 1 2">

(*H-Y« V bz ) *~n

^(iD+^i+Vj+^^f)

RH ,i(y-)forH20 ,

and

M--v« :R̂> : for oxveen

for carbon dioxide ,

bz 1 " " 2 J

The analysis of the surge bin is carried out in terms of finite time increments (At). The mass-balance equation for shale going from the mth to the nth increment is

( l - f K n At

( l - « ) At ^

Y(RD,j + RM,j)n (kg/m3-s)

and that for the gas is

GC02 ,n = ^ a j ( R D , j + ^M,j)n ' (kgC0 2 /m 3 - s )

where

At = l/50th of the surge bin residence time (s)

On the basis of the above-described mass balances and accounting for interphasial mass and heat transfer, we used the following equations in the combustor analysis.

The energy-balance equation for the solids phase (J/m3 • s) is

aiUiCPJ *2M£ '3a^\,T ^ L ~ ( * * i = h r ) ̂ - T0 +a* [-£) W r « - Ti) + Qi .

where

Q, = ^«j (R C j AH c + RHjAHH + Rs,jAH5 + RD/jAHD + RM,jAHM)

and

—- = mass of solid species transferred to the gas phase dz

The gas energy-balance equation is written as

iUfp* •&)+*& = -s(^)ws)

£ « , RCj + %, + RDJ + RM,j + *s{64) CP-i (T8 ~ T») 88

13

and

—- = mass of gas species transferred to the solid phase . dz

The energy balance for the surge bin is determined in a much different manner. The material entering the surge bin is assumed to have reached thermal equilibrium. Then, as carbonate decomposition occurs, the endothermic heat of reaction causes the solids temperature to drop. Hence, the energy balance need not account for heat transfer or individual particle temperatures. The overall energy-balance equation applied to the bin for time increment n, as time is incremented from m to n, is

( 1 -<KnCp,N,n(T n -T r ) ^ + °C0 2

L p-C0 2 Vn ~ *r) -

^ *'''s,nr-'p,N,m I i m * r ) V~< / n . „ , n , , , \ / i / 3 \

At ~ Z ai (RD,JAWD + V J A H M ) • (J/mJ • s)

Auxiliary Equipment

Auxiliary equipment is modeled to determine its effect on the system. This equipment includes blowers and heat-exchangers for the recycle gas to the retort and air to the combustor, as well as a conder ser for the product stream from the retort. Cyclones are also included, but only in the sense that they separate streams perfectly; no pressure drop or heat loss is calculated for these units. The power requirements for the blowers are calculated in two steps, assuming 80% efficiency. First the temperature change is calculated for the gas, assuming isentropic compression. Then the temperature (hence power requirement) is increased to account for the 20% inefficiency of the blower. The equation and appropriate constants are

T'-T{TJ • <K)

and the total temperature change across the compressor is

xrJhzIA , (K) 0.8

where

6 = ratio of the gas-specific heats Cp/Cv

= 1.4 for air = 1.3 for recycle gas .

The cooling load required to cool the gas and vapor products to 10 °C is calculated using the latent heats and specific heat equations listed in the Appendix. The assumption is that all water vapor is re­moved from the product stream so none is recycled. To account for latent heats, we estimate the fraction of oil products in the vapor phase (as opposed to condensed oil in mist form) from

_ (Xa - 273\ /ov \ 500 J '

14

where

Toi! = local gas-phase temperature . (K)

Steady-State Solution of System

The mode] arrives at a steady-state solution of the system by iterating through the reactor loop. For the first set of calculations, the retort is assumed to be isothermal, using an estimated temperature. In all subsequent calculations the previously calculated temperature for each respective stage is used. These retort calculations are repeated for each loop iteration until the temperature change between each iteration and the one that precedes it is less than 0.1 °C. Once this criterion has been met, combustor calculations begin, using the results of the retort calculations. When the combustor analysis is complete, the solids temperature and mass flow rate are used as input for the next set of retort calculations. This process continues until the desired level of accuracy is attained. The level of accuracy may be gaged by the convergence of combustor analysis results, such as mean solids temperature or height. When all the convergence criteria have been met, a system mass and energy balance is calculated.

The mass balance is evaluated from

Input (kg/s) == (raw shale -f air) flow rates.

Output (kg/s] = (burned shale + flue gas + oil -f gas products + foul water) flow rates.

And the energy balance is evaluated from

Input (kj/s) = (raw shale + air) enthalpies -+- exothermic heats of reaction + heater requirements + blower heat gains.

Output (kj/s) = (burned shale + flue gas + products) enthalpies + endothermic heats of reaction + condenser cooling requirements.

Application of the Model

The model was developed in a manner that would ease the carrying out of parametric studies. For both the retort and combustor, one can choose either a reactor characteristic and calculate the resulting condition, or vice versa. In simulating the retort, one can select a residence time and calculate the degree of kerogen pyrolysis that would result or can choose a desired conversion and calculate the required residence time. For the combustor simulation, it is possible to choose an outlet shale temperature and calculate the required height or to input the combustor height and calculate the corresponding exit tem­perature. These four system parameters, combined with the other input discussed previously, make the model practicable for extensive parametric studies.

Case Simulation

To show how the model may be applied, we did a hypothetical case simulation for a commercial operation :o produce an oil-plus-gas equivalent of 50,000 bbl/day. The raw shale properties and pertinent system conditions for this operation are given in Table 3. Note that two particle size distributions were input to the retort model and that the resulting mean particle diameter is listed in Table 3. The particle size distributions used in this case simulation were determined by J. F. Carley23 and are based on preliminary observations of attrition studies. The raw shale size distribution is meant to represent — 3 + 200 mesh shale.

15

Table 3. Conditions for sample-case simulation.

Raw shale composition

Grade (gal/ton) 30 Dolomite (wt%) 28.5 Calcite (wt%) 8.4 Bound water (wt%) 1.5

System conditions

Degree of kerogen pyrolysis (%) 98 Solids mass recycle ratio 3.5 Solids recycle temperature (°C) 665 Mean particle diameter in retort (mm) 0.7 Number of stages in retort 10

To run the case, we used an nth-order reaction model for kerogen pyrolysis9, with n = 1.4. The raw shale processing rate was set at 15000 kg/hr-m2 bed, with 4 wt% fines. Other input items not listed in Table 3 include the raw shale inlet temperature of 10 °C and retort fluidizing gas and combustor air inlet temperatures of 500°C. At the combustor, the air stoichiometric ratio was set at 1.0 and the exit pressure was set at 0.85 atm absolute.

Table 4 is a summary of the calculated results for this case simulation, and Fig. 2 is a flowsheet of the process results. The yield calculation is based on 100% Fischer assay being the maximum possible yield. The solids mixture temperature shown in the retort in Fig. 2 is the hypothetical equilibrium temperature that would be reached if no reactions were to take place; it is meant to illustrate the temperature drop occurring across the retort. Gas velocities above the retort bed and at the combustor exit were calculated as 1.0 and 14.0 m/s, respectively. Although the results of this simulation appear to be consistent and reason­able, they must be confirmed by additional experimental work.

Table 4. Summary of results of sample-case simulation.

Calculated results for retort

Residence time (min) 3.67 Heating value of dry product gas (kj/mol) 654 Superficial gas velocity at inlet (m/s) 0.33

Calculated results for combustor

Calcite decomposed (%) >0.1 Dolomite decomposed (%) 4.7 Superficial air velocity at inlet (m/s) 10.9

16

Gas products 10.8 kg/s -* ;;

10°C 0.80 atm

Recycle and products 131 kg/s

462°C 0.85 atm

Recycle blower 3.23 MW

-fi-^l

Raw shale 697 kg/s

10°C

29 kg/s 69.6°C

1.57 atm

Residence time 3.7 min

Mixture temp

•o (0

Q> © T3 O C CJ o o

Foul water 13.5 kg/s

Oil products 77.5 kg/s -*•

10°C

Solids recycle

2344 kg/ 665°C

Flue gas 328 kg/s

665°C 0.85 atm

Recycle heater

30.8 MW

500°C

Solids 2915 kg/s

503°C

Burned solids

576 kg/s 665°C

414.6 MW

Air 309 kg/s

500°C .888 atm Heater 156 MW

0.80 atm

Figure 2. Flowsheet of process results.

Acknowledgments

Helpful discussions with A. , Bumham, ,. , Carle.D. , Christiansen, A. B. Lewis, , C. Mallon,

H. Y Sohn and R. W. Taylor are gratefully acknowledged.

17

References 1. J. H. Campbell and A. K. Burnham, "Reaction Kinetics for Modeling Oil Shale Retorting," IN SITU 4,

1 (1980). 2. A. K. Burnham, "Reaction Kinetics and Diagnostics for Oil Shale Retorting," Inst, of Gas Tech., Symp.

Papers, Synthetic Fuels from Oil Shale II, Nashville, 1981 (Institute of Gas Technology, Nashville, TN, 1981).

3. J. H. Campbell, Modified ln-Situ Retorting: Results from LLNL Pilot Retorting Experiments, Lawrence Livermore National Laboratory, Livermore, CA, UCRL-53168 (1981).

4. R. L. Braun, Mathematical Modeling of Modified ln-Situ and Aboveground Oil Shale Retorting, Lawrence Livermore National Laboratory, Livermore, CA, UCRL-53119 (1981).

5. P W. Tamm, C. A. Bertelsen, G. M. Handel, B. G. Spars, and P. H. Wallman, "The Chevron STB Oil Shale Retort," Energy Progress 2, 37 (1982).

6. H. I. Weiss, "Retorting of Oil Shale: Background Status and Potential of the Lurgi Ruhrgas (LR) Process," in Proc. ACS Symp. Oil Shale Retorting—Latest Developments, Las Vegas, 1982 (American Chemical Society, Las Vegas, NV, 1982).

7. P. H. Wallman, P. W. Tamm, and B. G. Spars, "Oil Shale Retorting Kinetics," ACS 2nd Chem. Cong. North Amer. Cont., Las Vegas, 1980 (American Chemical Society, Las Vegas, NV, 1980).

8. J. H. Richardson, E. B. Huff, L. L. Ott, J. E. Clarkson, M. O. Bishop, J. R. Taylor, L. J. Gregory, and C. J. Morris, Fluidized-Bed Pyrolysis of Oil Shale: Oil Yield, Composition and Kinetics, Lawrence Livermore National Laboratory, Livermore, CA, UCID-19548 (1982).

9. R. L. Braun and A. K. Burnham, "Retort Modeling," in Oil Shale Project Quarterly Report, ]. F. Carley, Ed., Lawrence Livermore National Laboratory, Livermore, CA, UCID-16986-82-2 (1982).

10. J. H. Campbell, G. J. Koskinas, and N. P. Stout, The Kinetics of Decomposition of Colorado Oil Shale: 1. Oil Generation, Lawrence Livermore National Laboratory, Livermore, CA, UCRL-50289 (1976).

11. J F. Carley, Heat of Kerogen Decomposition and Improved Enthalpy-Temperature Relationships for Raw and Spent Colorado Oil Shales, Lawrence Livermore National Laboratory, Livermore, CA, UOPKK-75-28 (1975).

12. M. F. Singleton, G. J. Koskinas, A. K. Burnham, and J. H. Raley, Assay Products from Green River Oil Shale, Lawrence Livermore National Laboratory, Livermore, CA UCRL-53273 (1982).

13. D. E. Christiansen, Incipient Fluidization Characteristics of Crushed Oil Shale, Lawrence Livermore Na­tional Laboratory, Livermore, CA, UCRL-87053 (1982).

14. H. Y. Sohn and S. K. Kim, "Intrinsic Kinetics of the Reaction Between Oxygen and Carbonaceous Residue in Retorted Shale," Ind. Eng. Chem. Process Des. Dev. 10, 550 (1980).

15. R. W. Taylor, "Char Oxidation Rates," LLNL Oil Shale Project Monthly Highlights for December 1981, Lawrence Livermore National Laboratory, Livermore, CA, UOPKK 82-2 (1982).

16. H. Y. Sohn, "The Law of Additive Reaction Times in Fluid-Solid Reactions," Met. Trans. B 9B, 89 (1978).

17. R. G. Mallon and R. L. Braun, "Reactivity of Oil Shale Carbonaceous Residue with Oxygen and Carbon Dioxide," Col. School Mines Quart. 71, 309 (1976).

18. A. K. Burnham and R. W. Taylor, Occurrance and Reactions of Oil Shale Sulfur, Lawrence Livermore National Laboratory, Livermore, CA UCRL-87052 (1982).

19. X Quong, Vertical Pneumatic Conveying of Mixed-Particle-Sized Oil Shale, Lawrence Livermore Na­tional Laboratory, Livermore, CA UCRL-88524 (1983).

20. E. S. Pettyjohn and E. B. Christiansen, "Effect of Particle Shape on Free-Settling Rates of Isometric Particles," Chem. Eng. Prog. 44, 157 (1948).

21. D. Kunii and O. Levenspiel, Fluidization Engineering (R.E. Kreiger Publishing Co., Huntington, NY, 1977), pp. 388-391.

22. Ibid., Figure 19. 23. J. F. Carley, "Attrition of Spent Shales During Airveying and Cycloning," in Proc. Oil Shale Symp.,

16th, Golden, CO, 1983 (Colorado School of Mines, Golden, CO, 1983).

JAC/sb

18

Appendix

Physical and Thermodynamic Properties

Heat Capacity of Solids (J/kg-K) The heat capacity of raw and spent shale is the sum of the heat capacities of their organic and

inorganic mass portions within the shale. The equations for organic and inorganic material* are, respectively,

CpfC = 488.2 + 4.54 T

and

CpN = 552.64 + 0.9216 T .

Heat Capacity of Fluids (J/kgK)

The heat capacity equations for oil and liquid water are

Cp,oll = 488.2 + 4.54 T

and

CP,H2O = 4184 .

Heat Capacity of Gas Species (J/kg-K) The hecit capacities of the gas species are estimated from equations given by Hougen et al.t The heat-

capacity equations are:

cP,co2 = 6 0 2-7 8 +0.96422 T - 0.00032474 T2 ,

CpCO = 948.87 + 0.27062 T - 0.00004 T2 ,

CP,H2O = 1658.72 + 0.61365 T + 0.0000107 T2 ,

CpiH2 = 14415.7 - 0.40678 T + 0.00098727 T2 ,

Cp#CH = 837.85 + 4.8142 T - 0.0011715 T2 ,

C n r „ = 323.09 + 5.29296 T - 0.00152639 T2 ,

CpNj = 964.8 + 0.2076 T - 0.0000103 T2 ,

and

rn = 799.8 + 0.4141 T - 0.0001314 T2 .

* J. F. Carli?y, Heat of Kerogen Decomposition and Improved Enthalpy-Temperature Relationships for Raw and Spent Colorado Oil Shales, Lawrence Livermore National Laboratory, Livermore, CA, UOPKK-75-28 (1975).

t O. A. Hougen, K. M. Watson, and R. A. Ragatz, Chemical Process Principles 0ohn Wiley and Sons, Inc., New York, NY, 1954).

19

The heat capacity of the gas stream in the combustor is calculated as the mass average of the heat capacities of the species involved.

Latent Heats of Vaporization (J/kg)

The latent heats of vaporization or condensation (as required) for oil and water are, respectively,

AH v 0 = 230,000

and

AHvW = 2,260,000 .

Viscosities of Gas Species ( k g / m s )

The local viscosity of each gas specie is assumed to be a function of temperature alone and is calculated from the general equation

where the following data, extracted from nomograms,* apply:

Specie

co2 CO H 2 0 H2

CH4

CHX

"o x 103

0.0136 0.0169 0.0088 0.0083 0.0100 0.0075

n

0.87 0.67 1.113 0.67 0.76 0.94 .

The stream viscosity is then calculated from

YY.„.M1 / 2

/ l IT" I

"s =

where

A-J'j = molar weight of the rth gas specie ,

and

Y = mole fraction of the rth specie in the gas stream .

In this formulation, we have assumed that oil vapor makes a negligible contribution to the stream viscosity.

* J. H. Perry and C. H. Chilton, Chemical Engineer's Handbook (Mc-Graw Hill Book Co., New York, NY, 1973), pp. 3-212-213.

20

The gas viscosity in the combustor is estimated from one equation only. That equation is

i/air = 3.7X 10 - 7 T g0 6 8 5 .

Kerogen Mass Fraction (kg kerogen/kg raw shale)

The mass fraction of kerogen in the raw shale may be estimated from the Fischer assay, expressed in gallons per ton (GPT). The equation suggested by Singleton et al.* is

K = 0.00539=i(GPT + 1.026) .

Specific Gravity of Raw Shale (kg raw shale/m3)

An equation was written to calculate the specific gravity of raw shale, given the Fischer assay (in GPT) of that shale. The equation given by Smith t was meant to calculate the yield from a known specific gravity. Solving the equation for pr (the negative root is the correct root to use) yields

pr = 3263.3 (1198.24 + 126.25 GPT)1/2

0.063126

Gas —Solid Heat-Transfer Coefficient ( J / m 2 K s )

The heat-transfer coefficient used in the combustor analysis may be entered as an input parameter or, if desired, calculated internally. If it is not given in the input, the coefficient is calculated from*

Aj = ^ ( 2 + 0.6NPr1/3NReJ

1/2) , j

where fcg (the gas thermal conductivity) is calculated from

kS = i7'- (J/m-K.s)

and

NPr (the Prandtl number) is constant and equal to 0.70 .

* M. F. Singleton, G. J. Koskinas, A. K. Burnham, and J. H. Raley, Assay Products from Green River Oil Shale, Lawrence Livermore National Laboratory, Livermore, CA, UCRL-53273 (1982).

f J. W. Smith, "Specific Gravity—Oil Yield Relationships of Two Colorado Oil-Shale Cores," Ind. Eng. Chem. 48, 441 (1956). * D. Kunii and O. Levenspiel, Fluidization Engineering (R. E. Krieger Publishing Co., Huntington, NY, 1977), p. 210.

21