Group Title: 7th International Conference on Multiphase Flow - ICMF 2010 Proceedings
Title: 15.4.2 - Fluid Catalytic Cracking Process Intensification using a Rotating Fluidized Bed in a Static Geometry
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 Material Information
Title: 15.4.2 - Fluid Catalytic Cracking Process Intensification using a Rotating Fluidized Bed in a Static Geometry Fluidized and Circulating Fluidized Beds
Series Title: 7th International Conference on Multiphase Flow - ICMF 2010 Proceedings
Physical Description: Conference Papers
Creator: Trujillo, W.R.
De Wilde, J.
Publisher: International Conference on Multiphase Flow (ICMF)
Publication Date: June 4, 2010
 Subjects
Subject: CFD
fluid catalytic cracking
rotating fluidized bed
 Notes
Abstract: The intrinsic potential of rotating fluidized beds in a static geometry for intensifying the fluid catalytic cracking process is evaluated by means of CFD simulations using an Eulerian-Eulerian model and the Kinetic Theory of Granular Flow. The reactions are described by a 10-lump model. The reaction kinetics is based on currently allowable cracking temperature and catalyst activity and a comparison with riser technology is made.
General Note: The International Conference on Multiphase Flow (ICMF) first was held in Tsukuba, Japan in 1991 and the second ICMF took place in Kyoto, Japan in 1995. During this conference, it was decided to establish an International Governing Board which oversees the major aspects of the conference and makes decisions about future conference locations. Due to the great importance of the field, it was furthermore decided to hold the conference every three years successively in Asia including Australia, Europe including Africa, Russia and the Near East and America. Hence, ICMF 1998 was held in Lyon, France, ICMF 2001 in New Orleans, USA, ICMF 2004 in Yokohama, Japan, and ICMF 2007 in Leipzig, Germany. ICMF-2010 is devoted to all aspects of Multiphase Flow. Researchers from all over the world gathered in order to introduce their recent advances in the field and thereby promote the exchange of new ideas, results and techniques. The conference is a key event in Multiphase Flow and supports the advancement of science in this very important field. The major research topics relevant for the conference are as follows: Bio-Fluid Dynamics; Boiling; Bubbly Flows; Cavitation; Colloidal and Suspension Dynamics; Collision, Agglomeration and Breakup; Computational Techniques for Multiphase Flows; Droplet Flows; Environmental and Geophysical Flows; Experimental Methods for Multiphase Flows; Fluidized and Circulating Fluidized Beds; Fluid Structure Interactions; Granular Media; Industrial Applications; Instabilities; Interfacial Flows; Micro and Nano-Scale Multiphase Flows; Microgravity in Two-Phase Flow; Multiphase Flows with Heat and Mass Transfer; Non-Newtonian Multiphase Flows; Particle-Laden Flows; Particle, Bubble and Drop Dynamics; Reactive Multiphase Flows
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Volume ID: VID00377
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Holding Location: University of Florida
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Resource Identifier: 1542-Trujillo-ICMF-2010.pdf

Full Text



7th International Conference on Multiphase Flow,
ICMF 2010, Tampa, FL, May 30 June 4, 2010


Fluid Catalytic Cracking Process Intensification using a Rotating Fluidized Bed in
a Static Geometry


Waldo Rosales Trujillo and Juray De Wilde

Unit6 IMAP, Universit6 catholique de Louvain, Place Sainte Barbe 2, 1348 Louvain-la-neuve, Belgium
waldo.rosales ~uclouvain.be and j uray.dewilde ~uclouvain.be

Keywords: CFD, Fluid Catalytic Cracking, Rotating Fluidized Bed




Abstract

The intrinsic potential of rotating fluidized beds in a static geometry for intensifying the fluid catalytic cracking
process is evaluated by means of CFD simulations using an Eulerian-Eulerian model and the Kinetic Theory of
Granular Flow. The reactions are described by a 10-lump model. The reaction kinetics is based on currently allowable
cracking temperature and catalyst activity and a comparison with riser technology is made.


Nomenclatu re


Roman symbols
Aij pre-exponential factor of the reaction of
formation of lump i out of lump j, (kmol/(m3 s))
C, molar concentration of lump j, (kmol/m )
Egy activation energy of the reaction of
formation of lump i out of lump j, (kJ/kmol)
e,, solid-solid restitution coefficient, (-)
d, particle diameter, (m)
go,ss radial distribution function, (-)
h bed height, (m)
I2D SOCOnd invariant of the deviatoric stress tensor,
(1/s )
K aromatic adsorption constant, (-)
rj2 mass flow rate, (kg/s)
Mi molar mass of lump i, (kg/kmol)
N Average number of catalyst turns in the
RFB-SG, (-)
p, ps gas phase pressure, solid phase pressure, (Pa)
Rij reaction rate of lump i out of lump j,
(kmol/(m3 s))
Eccoking rate, (kmol/(m s))
R' coking rate per mass catalyst, (kgcoke/(kgats))
Sc mass source term in the continuity equation,
(kg/(m3 s))
S, source term of lump i in the species continuity
equation, (kg/(m s))
T temperature, (K)
it, it gas phase and solid phase velocity, (m/s)
V reactor volume, (m )
wi mass fraction of lump i, (-)


Greek symbols
Sinterphase momentum transfer coefficient,
(N/m )
Ye Collisional dissipation of energy, (J/(s m ))
phase volume fraction, (-)
8 granular temperature, (J/kg)
ec~ granular diffusion coefficient, (-)
Viscosity, (Pa s)
vik stoichiometric coefficient ratio of the
formation of lump i out of lump k, (-)
p,, ps gas phase density, solid phase particle density,
(kg/m )
ret catalyst residence time, (s)
-r phase stress-strain tensor, (N/m )
angle of internal friction, (rad)
XGo Gas Oil conversion, (-)
Subscripts
cat catalyst
g, s gas phase, solid phase
GO, C Gas Oil, C lump,
G, LG gasoline, Light Gases
h, I heavy lump, light lump


Introduction

Rotating fluidized beds in a static geometry (RFB-SG)
have been recently developed (De Wilde & de Bro-
queville 2007, 2008; de Broqueville & De Wilde 2009).
The rotational motion of the particle bed is introduced
by the tangential injection of the fluidization gas in the







7th International Conference on Multiphase Flow,
ICMF 2010, Tampa, FL, May 30 June 4, 2010


ial gas velocities in conventional fluidized beds or risers
(De Wilde & de Broqueville 2007, 2008; de Broqueville
& De Wilde 2009). This combination causes the gas
phase residence time in RFB-SGs to be easily one to
two orders of magnitude smaller than in conventional
fluidized beds or in risers. Looking at reactor applica-
tions of RFB-SGs, the small gas phase residence time
may be limiting for the conversion. To compensate for
the small gas phase residence time, RFB-SGs offer parti-
cle bed densities significantly higher than those in risers.
Furthermore, High-G fluidization was shown to improve
the particle bed uniformity, bubbling being suppressed.
As such, bypassing of solids (catalyst) by the gas, en-
countered in conventional fluidized beds, can be avoided
in RFB-SGs and the gas-solid contact improved. Also,
the gas-solid slip velocity can be much higher than in
COnventional fluidized beds or risers, which is advanta-
geous for gas-solid mass and heat transfer.
To evaluate the process illiaislilk~.lil nI potential of
RFB-SG type reactors accounting for the impact of the
different above mentioned factors, CFD simulations of a
RFB-SG used for fluid catalytic cracking (FCC) are car-
ried out. The intrinsic process illiaislilk~.lil nI potential
of RFB-SGs is evaluated, that is, using the same cata-
lyst and kinetics used in riser simulations (Froment &
Bischoff 1990) and a typical cracking temperature of
775K.


Simulation Model

Continuity Equations The Eulerian-Eulerian two-
phase model considers two fully penetrating and inter-
acting phases. Mass, momentum and species continuity
equations are solved for each phase (Table 1). The gas
phase density variations due to cracking have to be ac-
counted for. The Kinetic Theory of Granular Flow was
used for the calculation of the solid phase physical prop-
erties (Gidaspow 1994). This requires the solution of an
additional continuity equation for the granular tempera-
ture, also shown in Table 1. The resulting constitutive
equations are shown in Table 2. An important model pa-
rameter is the restitution coefficient for particle-particle
collisions, e ss, which was given a value of 0.9.
Both inter-phase momentum and mass transfer are ac-
counted for. The latter results from coke deposition on
the catalyst. The drag coefficient model of Gidaspow
(1994) was used for the calculation of the gas-solid mo-
mentum transfer. Gas-solid species mass and heat trans-
fer rate limitations can be neglected due to the high gas-
solid slip velocities and related gas-solid mass and heat
transfer coefficient (de Broqueville & De Wilde 2009).
Particle bed mixing and heat transfer in the rotating par-
ticle bed were shown to be extremely fast (De Wilde &
de Broqueville 2007). If the time scales of particle bed


nt g


soid solids


Figure 1: Schematic representation of the Rotating Flu-
idized Bed in a Static Geometry Reactor
(RFB-SG).


static fluidization chamber via multiple slots (Figure 1).
Both tangential and radial fluidization of the particle bed
can be obtained. Radially, the centrifugal force is coun-
teracted by the gas-solid drag force. The former can be
multiple times gravity, allowing a higher gas-solid drag
force and as such overcoming the limitations of conven-
tional fluidized beds. In particular, operation with small
particle bed heights and high panticle bed densities at
high gas velocities is possible.
High-G operation was also shown to allow the
fluidization of smaller, 1G-Geldart-C type particles
(Watano et. al. 2004). An important characteristic of
RFB-SGs is the flexibility in the fluidization gas flow
rate, the latter affecting the centrifugal force and the
counteracting radial gas-solid drag force in a similar way
(de Broqueville & De Wilde 2009; De Wilde & de Bro-
queville 2008).
De Wilde & de Broqueville (2007) presented experi-
mental proof of concept of RFB-SGs and studied their
basic hydrodynamics, stability, and mass and heat trans-
fer characteristics (de Broqueville & De Wilde 2009).
The excellent particle bed mixing property of RFB-SGs
was experimentally and numerically demonstrated (De
Wilde & de Broqueville 2008; de Broqueville & De
Wilde 2009). It results, for example, in significantly im-
proved panticle bed temperature uniformity.
Whereas in conventional fluidized beds or risers, the
particle bed height is of the order of meters or tenths
of meters, in RFB-SGs, the particle bed height is only
of the order of centimeters. Funthermore, the radial gas
velocities in RFB-SGs can be much higher than the ax-





7th International Conference on Multiphase Flow,
ICMF 2010, Tampa, FL, May 30 June 4, 2010


8t(E pg) + V -(E ,* .7)


Mass




Species




Momentum




Granular Tem-
perature
Source Terms


-Scoke


~(tsps) + V (Isps&;) =Scoke

8t(espwsi) V (e pglw ) = S% (for lump i)

8(EsPsacoke) E.(sPs6 coke) = Scoke


i3
(f , ;1)+~(f ;11)
i3t
i3
(~sps~+ V(~,pv'v~
i3t


:9Vp + V 79 + P(6U


,o~? o~?, + o 7, + P(UJ


ji[(p.-" sE^) Ci- )=(-pl rs) C: V&+ V-nl ( VS T


Scoke = coke Mcoke Ok~lVCke


Table 1: Continuity Equations.


mixing and heat transfer are assumed to be small com-
pared to the coking time scale and the related average
residence time of the particles in the reactor, the tem-
perature and catalyst coke content in the reactor can be
assumed uniform and can be imposed in the simulations.
If, on the other hand, the catalyst residence time is low
to limit the increase in catalyst coke content, a uniform
coke content assumption is also justified.
A Reynolds-Averaged approach was taken and turbu-
lence was accounted for using the k -t turbulence model
for each phase. A grid-independent solution was guar-
anteed and the grid was sufficiently refined to calculate
possible meso-scale phenomena, like bubbles. The latter
implies a non-stationary (transient) simulation. The grid
size was as small as 0.2 mm and the time step 1 x 10-s
s. Calculations were continued until a statistically sta-
tionary state was reached.
At solid walls, a no-slip condition for the gas phase
and a partial slip condition for the solid phase were im-
posed. The Johnson-Jackson model (Johnson & Jackson
1987) was adopted, using a specularity coefficient of 0.2
and a particle-wall restitution coefficient of 0.9.


Reaction Mechanism and Kinetics The catalytic
cracking of Gas Oil was described using the ten-lump
model of Jacob et. al. (1976). A scheme showing the
lumps, as well as the possible reactions, is shown in Fig-
ure 2.
The reaction rate for the formation of lump i out of


Figure 2: Ten-Lump model for the catalytic cracking of
Gas Oil.Jacob et. al. (1976)



lump j is calculated from:


A, -e
Rij = -ps~s C3
1+Kh ,W ,


The molar masses of the different lumps as well as
the pre-exponential factors and activation energies of
the different reactions can be found in Jacob et. al.
(1976). The effect of the adsorption of heavy aro-
matics is accounted for via the non-dimensional term
(1 + Ka, Wp, ) Catalyst deactivation is modeled
as a function of the coke content on the catalyst by





7th International Conference on Multiphase Flow,
ICMF 2010, Tampa, FL, May 30 June 4, 2010


p, = 2ps (1 e 2 gns


Solid phase pressure

Radial distribution


Solid phase collisional viscosity

Solid phase kinetic viscosity


Solid phase frictional viscosity


Gas-solid drag force


go ss = +



V ~coin = -- p lso (1 ss
6(3 ~ ~ e,

p, si =


~,< 0.2 : ,3 = -CD 99-2t


('D 24[1 + 0.75( ,Re, )n.68s
ysRe,


Table 2: Constitutive Equations. Gidaspow (1994); Syamlal et. al. (1993); Syamlal (1987)


o (1 + 69.47 1001oit.j'coe3
The (' lump is considered as a mixture of (solid
phase) coke and light gases (C1-C4). A fixed value
fooke of 0.19 was used. For the calculation of the phys-
ical properties of the lump of Light Gases, propane was
taken as the representative component.
The code describing the kinetic model was validated
by reproducing the 1D riser simulations of Froment &
Bischoff (1990).


Simulation Cases and Set-up

The problem is simulated in 2D using FLUENT
v.3.6.26. User defined functions describe the reaction
kinetics. The calculation domain represents a periodical
section of a 12-slot polygonal body reactor and includes
a section of the gas distribution chamber. The latter al-
lows for pressure changes in the particle bed to be trans-
ponted upstream the injection slot. Catalyst is continu-
ously and nearly tangentially fed via a side inlet. When
reaching the statistically stationary state, catalyst losses
via the chimney compensate for the continuous feeding
of catalyst. A separate catalyst outlet in the dense par-
ticle bed region could be considered as, for example, in
the experiments by De Wilde & de Broqueville (2007).
The reactor dimensions and operating conditions are
summarized in Table 3. Gas Oil conversion, gasoline
and Light Gasses yield and selectivity are compared with
values obtained using a 30m tall riser reactor, taken from
Froment & Bischoff (1990), and the intrinsic process in-
tcellsilk~.ll ill potential is discussed.
Convergence was assumed to be reached when the


Gas Inlet


Outlet


Figure 3:Periodic calculation domain of a 12-slot
polygonal body reactor.





residuals dropped below le-6 for the energy equation
and le-3 for the other continuity equations. Funther-
more the overall mass balances were verified. Typically,
reaching a statistically steady state required 60s, corre-
sponding to 600h of calculation time on a single proces-
sor.


1) s ynes s













Distribution Number gas of 12-
chamber inlets
Gas inlets width 15 mm
Gas velocity at 19.6 m/s
the inlets
Gas density 4.72 kg/m

Reactor Number of slots 12 -
Slots width 3 mm
Number of cata- 12 -
lyst inlets
Catalyst inlets 2.6 mm
width
Reactor outer ra- 0.6 m
dius
Reactor inner ra- 0.2 m
dius (chimney)
Particle diameter 80 m
Particle density 1500 kg/m3
Catalyst velocity 15 m/s
at the inlets
Catalyst volume 0.3 -
fraction at the in-
lets
Catalyst loading 23 kg/m
Catalyst coke 0.15 wt-%
content

GO Lump ag(%) AL
Pi, 17 339
N;, 20 339
Asi, 24 339
Ari, 12 339
P1 14 226
N1 9 226
Asi '> ''6
Art '> ''6
G 0 114
C 0 34


Table 3: Reactor dimensions and operating conditions.
Gas Oil (GO) composition and molar mass of
the different lumps. From Jacob et. al. (1976).



Results and Discussion


Figure 4 shows the contour plot of the solids volume
fraction in the RFB-SG. Except in the immediate vicin-
ity of the outer reactor wall, where the effect of the
polygonal design and a limited non-uniformity due to
the gas and solids injection are visible, the tangential
uniformity of the solids volume fraction profile is ac-
ceptable. A radial profile in the range 0.2m the tangentially averaged solids volume fraction in the


7th International Conference on Multiphase Flow,
ICMF 2010, Tampa, FL, May 30 June 4, 2010


3.00.-01
2.88e-01
2.70.-01
2.4e-01
2.40e-01
2.28e-01
2.16e-01
2.04e-01
1.92.01
1.80e-01
1.68.-01
1.56e-01
1.44e-01
1.32e-01
1.20e-01
1.ase-ai
9.60.-02
8.40e-02
7.20e-02
6.00e-02
4.eae-02
3.60e-02
2.40e-02
1.20e-02
O.coe+oo


Figure 4: Contour plot of the solids volume fraction in
the reactor. Reactor characteristics and oper-
ating conditions: see 3.


RFB-SG is shown on Figure 5 in which a comparison is
made with the axial profile of the cross-sectional aver-
aged solids volume fraction in the riser. The coordinates
were normalized for comparison purposes. In the RFB-
SG, the loss of a well separated particle bed freeboard
is resulting both from the density variations due to reac-
tions and from the continuous solids feeding required to
maintain a sufficiently high solids loading in the reactor.
Values for the solids volume fraction in the RFB-SG
range from 0.15 to 0.28, with an average of 0.225, which
is about 7.5 times higher than the average value in the
riser. In the RFB-SG, the low solids volume fraction
in the immediate vicinity of the outer wall is caused by
the gas injection and was confirmed experimentally (De
Wilde & de Broqueville 2007, 2008). Bubbling does not
occur and a stationary, rather than statistically stationary
state is achieved.
In Figure 6, radial profiles of the tangentially averaged
tangential and radial solid phase velocities are shown.
The average radial gas velocity is 2.67 m/s, implying a
gas phase residence time of 0.15 s. The mean particle
bed rotation speed is 15 m/s, corresponding to a cen-
trifugal acceleration of approximately 38G at the outer
wall of the reactor. As such, the average solids rotation
period is 0.19 s.
7 shows the Gas Oil and gasoline mass fraction con-
tours in the RFB-SG reactor. The Gas Oil mass fraction
is calculated as the sum of the mass fractions of the dif-
ferent lumps of which Gas Oil is composed. The Gas
Oil conversion XGO can be defined as 1 wGO, the con-












wi(%) Ph Nh Ash Arh P1 N1 As1 Arl G LG coke XGO %b)
RFB-SG 10.12 9.41 7.73 10.32 12.90 7.24 3.26 6.88 23.38 7.09 1.67 32.14
riser 6.14 4.02 1.89 8.7 11.77 5.37 3.51 9.23 37.54 9.58 2.25 49.36

Table 4: Lump mass fractions and GO conversion comparison between RFB-SG and the riser reactor technology.
RFB-SG reactor characteristics and operating conditions: see 3. RFB performance with conventional catalyst
and cracking temperature of 775K. Riser data from Froment & Bischoff (1990).


7th International Conference on Multiphase Flow,
ICMF 2010, Tampa, FL, May 30 June 4, 2010


-(a) RFB-SG

--- (b) Riser


O


v,0.2


0 0.2 0.4 0.6 0.8
Normalized Coordinate (z/L)


0 0.1 0.2 0.3 0.4 0.5
Normalized coordinate (z/L)


Figure 5: Radial profile of the tangentially averaged
solids volume fraction in the RFB-SG and in
a riser reactor. Reactor characteristics and op-
erating conditions: see 3. Riser data from
from Froment and Bischoff (1990)Froment &
Bischoff (1990)


version into gasoline as XG = G, and that of Light
Gases aS XLG = LG. A maximum Gas Oil conversion
of 31.24% and conversion into gasoline of 23.38% is ob-
tained in the RFB-SG.
Figure 4 compares the mass fraction of the differ-
ent lumps and the Gas Oil conversion obtained at the
outlet of the RFB-SG and the riser reactor (Froment &
Bischoff 1990). A comparison between the Gas Oil and
gasoline conversion profiles in the RFB-SG and in the
riser is shown in Figure 8. For the given RFB-SG ge-
ometry and operating conditions, results indicate some-
what lower conversions compared to the riser. Essential
is, however, the Process ]Illiais~ilk~.Iil 'II factor, which ac-
counts for the differences in Gas Oil flow rate and reac-
tor volume, as discussed hereafter. If required, the con-
version in the RFB-SG can be increased by increasing
the gas phase residence time (increasing the particle bed
height or reducing the gas flow rate), or by increasing
the reaction rates (increasing the operating temperature
or catalyst activity). The latter two options will be con-
sidered in more detail further in this paper.
The process illiaisilk~.Ilisl I (PI) potential can be de-
fined as the ratio of the amount of gasoline produced per


Figure 8: Gas Oil conversion profile and conversion
into gasoline profile. Comparison between
RFB-SG and riser type reactor. RFB-SG char-
acteristics and operating conditions: see Table
3. Riser data from (Froment & Bischoff 1990)


unit time and per unit volume reactor (2).


XRFB~Xo
PI(xco) = G where Xc(Xco)
"; co)


wo(xoo) .mGO
V Xc)


In Equation (2), wG XGo) is the gasoline mass frac-
tion produced for given Gas Oil feed rate, rj2GO, and Gas
Oil conversion XGo. V(XGo) is the particle bed vol-
ume required to achieve this Gas Oil conversion at the
the given Gas Oil feed rate.
The PI factor was calculated as a function of the Gas
Oil conversion,XGcO, and is shown in Figure 9. The PI
factor increases with the Gas Oil conversion and values
as high as 7 can be easily achieved. This is mainly due to
the higher particle bed density and the improved particle
bed uniformity.
Finally, the high cat-to-oil ratios at which RFB-SGs
can be operated should be remarked. Typical values in
riser reactors are of the order of 10, maximum 20 (Fro-
ment & Bischoff 1990). In the RFB-SG simulated, the
gas and catalyst motion is mainly cross-flow and the cat-
to-oil ratio, calculated from Equation (3), amounts to
roughly 500.









7th International Conference on Multiphase Flow,
ICMF 2010, Tampa, FL, May 30 June 4, 2010


9.60e404
9.20e+04

11.20e045
8.80e+04
8.40e+04
I8.00e+04
7.60e+04
7 20e+04
6 80e404
6.40e+04
6.00e+04
5.60e+04
5.20e+04
4.80e+04
4.40e+04
4.00e+04
3.60e+04
3.20e+04
2.80e+04

2.00e404
1.60e404
1.20e+04
8.00e403
4.00e+03
0.00e+00


4.73e+00
4.66e+00
4.59e+00
4.53e+00
S4.46e+00
4.39e+00
4.33e+00
4.26e+00
4.19e+00
4.13e+00
4.06e+00
3.99e+00
3.93e+00
3.86e+00
3.79e+00
3.73e+00
3.66e+00
3.59e+00
3.53e+00
3.46e+00
3.39e+00
3.33e+00
3.26e+00
3.19e+00
3.13e+00
3.06e+00


(a) Static pressure (in Pa)


(b) Fluid density ...~ I .


30


S25


'O20


S15





6 0



0.6 0.5 0.4 0.3

Radial Coordinate [rn]


0.8

0.6 2

0.4

0.2


--Tangential
--- Radial


Figure 6: Radial profiles of the tangentially averaged tangential and radial solid phase velocities in the RFB-SG.
Reactor characteristics and operating conditions: see Table 3


A contour plot of the calculated coking rate in the
RFB-SG is shown Figure 10, an average value is shown
in Table ??, and allows an average particle residence
time of 0.81 s. This implies that the catalyst is to be

replaced on average every 4.03 turns, which requires a
catalyst feeding or regeneration rate of 181.4 kg/s per
meter of reactor axial length. In the simulation, the cata-

lyst feeding rate was 702 kg/s per meter of reactor axial
length.

An alternative approach for facilitating the treatment
of the high solids mass flow rates and to be considered
in more detail in future work is the combined catalytic
cracking and catalyst regeneration, that is, in a single
vessel, using a multi-zone concept type RFB-SG reac-
tor. A possible multi-zone concept is shown in Figure
11. Gas Oil is injected in the cracking zone, whereas


rgPg (Urad)


cat-to-oil ratio


The maximum allowable average solids residence
time can be estimated in terms of the allowable coke

content variation, at the given coking rate profile, 8tc.
If an average 5% coke content increase is admitted:




7cat = .5ck'Pwhere SRcoke /coke "Gck~cc (4)
ncoke coRke)


The required catalyst feed or regeneration rate can
then be calculated from:

esps Vbed
me~at = Tct(5)


'










7th International Conference on Multiphase Flow,
ICMF 2010, Tampa, FL, May 30 June 4, 2010





2.38e-01
2.28e-01
2.19e-01
2.09e-01
2.00e-01
1 90e-01
1 81e-01
1.71e-01
1.62e-01
1.52e-01
1.43e-01
1.33e-01
1.24e-01
1.14e-01
1 05e-01
9 51e-02
8.56e-02
7.61e-02
6.66e-02
5.71e-02
4.76e-02
3.80e-02
2.85e-02
1.90e-02
9.51e-03
000e+00


9.88e-01
9.75e-01
9.63e-01
9.50e-01
S9.38e-01 .0e0
9.26e-01
9.13e-01
9.01e-01
8.88e-01
8.76e-01
8.64e-01
8.51e-01
8.39e-01
8.26e-01
8.14e-01
8.02e-01
7.89e-01
7.77e-01

7.52e-01
7.40e-01
7.27e-01
7 15e-01
7 03e-01
S 76F0690e-01


(a) Gas Oil mass fraction


(b) Gasoline mass fraction


6.92e-02
6.63e-02
6.34e-02
6.05e-02

5.48e-02
5.19e-02
4.90e-02
4.61e-02
4.32e-02
4.04e-02
375e-02
346e-02
3 17e-02
2.88e-02
2.59e-02
2.31e-02
2.02e-02

1.44e-02
1.15e-02
8.65e-03
5.77e-03
2.88e-03
6.00e-09


0.35

0.3 *

0.25

O.-'
0.2 ,

0.15



0.05 ,f-i

0
0.60 0.50 0.40 0.30

Radial coordinate (m)


--- Gerio


(c) Light Gases mass fraction




Figure 7: Contour plots of the mass fraction of (a) Gas
conditions: see Table 3.




air is injected in the catalyst regeneration zone. The two

zones could be separated from each other by additional

steam injected 'buffer' zones. The multi-zone concept

also opens perspectives for increasing the efficiency of

heat transfer between the catalyst regeneration and the

catalytic cracking zones, allowing further intensification
of the FCC process.




Conclusions



The potential of rotating fluidized beds in static geome-

try for the illiaisilk~.Iion of the fluid catalytic cracking

process has been confirmed by CFD simulations using
the Eulerian-Eulerian approach with the Kinetic The-


(d) Gas Oil conversion and conversion into gasoline and Light
Gases



Oil and (b) gasoline. Reactor characteristics and operating






ory of Granular Flow and a 10-lump reaction model. A
one order of magnitude intrinsic process illiaislilk~.Il il a

can be easily achieved, that is, using the conventional

cracking catalyst and temperature. The gasoline or Light

Gases selectivity can also be optimized more easily.








The improved particle bed mixing and related unifor-

mity may also allow the use of higher cracking temper-
atures and/or a more active catalyst and further intensi-

fication of the FCC process. This should be taken into

consideration when evaluating the full process intensifi-
cation potential of RFB-SGs.








7th International Conference on Multiphase Flow,
ICMF 2010, Tampa, FL, May 30 June 4, 2010


4.20e-02
4.03e-02
3.85e-02
3.68e-02
I .803.50e-02
3.33e-02
3.15e-02
2.98e-02
2.80e-02
2.63e-02
2.45e-02
2.28e-02
2.10e-02
1.93e-02
1.75e-02
1.58e-02
1.40e-02
1.23e-02
8.76e-03
7.01e-03
5.26e-03
3.50e-03
0.000+00


10% 20%
GO Conversion


Figure 10l: Catalyst coking rate (in kmol/m s) in the
RFB-SG. RFB-SG reactor characteristics
and operating conditions: see 3.











Fiue 1:Tneta ut-oecnet neto fd


feen eoe hog h ifeetgsiltsos

1aan 50e-01a ua .ad a aa .ad aa
1as 38e-01ae .ad aeR.N ndPefe .
1in 27e-01 otn b oelrttigfudie e
coatr, owdr Tehnoogy Vol 14, 3 pp.17217'










Figuure 11angiential muti-zone concpt.Ionje ctinfdeif-


Watano S. nd Nakmur H.ad aaa ad aa

Fie parile co atngb ae nroquvlel Arotating fluidizedbe

coate, PIh owdr Technology3,4,Vol. 14-1, p.12-176


Froen G.d F. and deBiscoff K. Be., Chpermical Reator



tff iinuum and Kinetic Theor D-yes aritiones, Academict


Figure 9: Intrinsic Process ]Illiais~ilk~.nll I potential of
RFB-SG for the catalytic cracking of Gas
Oil as a function of the Gas Oil conversion.
Comparison with riser technology (Froment
& Bischoff 1990). RFB-SG reactor charac-
teristics and operating conditions: see Table




Acknowledgements

The authors thank the "Fonds de la Recherche Scien-
tifique (FNRS)" for the financial support, under project
FRFC no 75394/2.4.597.07 F. The authors also acknowl-
edge the "Institut de calcul intensif et de stockage de
masse (CISM)" for the parallel computing infrastructure
and technical support.


References


Jacob S. M. and Gross B. and Voltz S. E. and Week-
man V. W. Jr, A lumping and reaction scheme for cat-
alytic cracking, AIChE Journal Vol.22 Issue 4 pp.701_
713, Wiley-Interscience, New York, 1976

Syamlal M., The Particle-Particle Drag Term in a Multi-
particle Model of Fluidization, National Technical Infor-
mation Service, Springfield, VA, DOE/MC/21353-2373,
NTIS/DE87006500, 1987

Syamlal M. and Rogers W. and O'Brien T. J,
MFIX Documentation: Volume 1, Theory Guide, Na-
tional Technical Information Service, Springfield, VA,
DOE/METC-9411004, NTIS/DE9400087, 1993

Johnson P. C. and Jackson R., Frictional-Collisional
Constitutive Relations for Granular Materials, with Ap-
plication to Plane Shearing, Journal of Fluid Mechanics
Digital Archive, Vol. 176 pp. 67-93, 1987







7th International Conference on Multiphase Flow,
ICMF 2010, Tampa, FL, May 30 June 4, 2010


ing fluidized bed with a rotating chimney, AIChE Jour-
nal, Vol. 54, 8, pp. 2029-2044, 2008

de Broqueville A. and De Wilde J., Numerical investiga-
tion of gas-solid heat transfer in rotating fluidized beds
in a static geometry, Chemical Engineering Science, Vol.
64, 6, pp. 1232-1248, 2009




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