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IJERA (International journal of Engineering Research and Applications) is International online, ... peer reviewed journal. For more detail or submit your article, please visit www.ijera.com

IJERA (International journal of Engineering Research and Applications) is International online, ... peer reviewed journal. For more detail or submit your article, please visit www.ijera.com

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  • 1. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714 Modelling And Simulation Of Industrial FCC Unit: Analysis Based On Five-Lump Kinetic Scheme For Gas-Oil Cracking K. K. Dagde * and Y. T. Puyate***(Department of Chemical/Petrochemical Engineering, Rivers State University of Science and Technology, Port Harcourt, P. M. B. 5080, Port Harcourt, Nigeria) **(Department of Chemical/Petrochemical Engineering, Rivers State University of Science and Technology, Port Harcourt, P. M. B. 5080, Port Harcourt, Nigeria)Abstract Models which describe the performance related heavy stocks, into a broad spectrum ofof the riser and regenerator reactors of fluid products in the presence of a catalyst. The productscatalytic cracking (FCC) unit are presented. The of catalytic cracking include fuel gases, liquefied petroleum gas, high-octane gasoline, light fuel oil,riser-reactor is modelled as a plug-flow reactor diesel fuel, heavy fuel oil, etc. FCC unit consists of aoperating adiabatically, using five-lump kinetics reaction section and a fractionating section thatfor the cracking reactions. The regenerator- operates together as an integrated process unit. Thereactor is divided into a dilute region and a dense reaction section has two reactors: (i) the riser-reactorregion, with the dense region divided into a where almost all the endothermic cracking reactionsbubble-phase and an emulsion phase. The bubble- and coke deposition on the catalyst occur, and (ii) thephase of the regenerator is modelled as a plug- regenerator-reactor, where air is used to burn-off the accumulated coke on the catalyst. The catalyst-flow reactor, while the emulsion phase is modelled regeneration process also provides the heat requiredas a continuous stirred tank reactor. The models for the endothermic cracking reactions in the riser-are validated using plant data obtained from a reactor.functional industrial FCC unit. It is shown that In the FCC unit, the catalyst enters the riser-predictions of the models compare very well with reactor as a dense bed and is pneumatically conveyedplant data for both reactors. Simulation results upwards by the dispersing steam and vapourizing gas-oil feed. It is during this period of conveying theindicate that catalyst-to-gas oil ratio and inlet-air catalyst that catalytic cracking of gas-oil takes placevelocity have significant effects on the through efficient catalyst and gas-oil contact. Theperformance of the riser and regenerator reactors catalyst later becomes deactivated due to cokerespectively. The yield of gasoline and other deposition on it, and the deactivated catalyst thenproducts of the catalytic cracking process increase passes through the spent-catalyst slide-valve in theas the height of riser-reactor increases, with riser-reactor and enters the top of the regenerator.maximum yield of gasoline (of about 0.45 mass The major purpose of the regenerator is to oxidize the coke on the spent catalyst with oxygen to form CO,fraction) occurring about half-way up the riser- CO2, and H2O, thereby reactivating the catalyst.height. Both the amount of coke on spent catalyst Compressed combustion-air enters the regeneratorand the riser-temperature decrease with time, from the bottom through a grid distribution patternwhile the regenerator-temperature increases with designed to provide efficient mixing of air withtime. The riser-temperature varies from about deactivated catalyst, resulting in a fluidized-bed650K to 800K, while the regenerator-temperature catalyst-regeneration operation. The regeneratedranges from about 650K to 1080K. The optimum catalyst passes through the regenerated-catalyst slide- valve and is mixed with gas-oil at the riser-reactor‟svalues of process variables obtained for effective base and the cycle is repeated. Provisions are madeoperation of FCC are inlet-air velocity of 14m/s, for adding fresh catalyst makeup to maintainriser-temperature of about 653K, and catalyst-gas inventory and for withdrawal of aged andoil ratio of 3. contaminated catalyst. The FCC unit is quite complex from both the process modelling/simulation andKeywords: Riser, regenerator, kinetics, catalytic control points of view. The complexity of the FCCcracking, simulation. unit is attributed to the strong interaction between the riser and the regenerator reactors, and the uncertainty1. Introduction in the cracking reactions, coke deposition, and coke Fluid catalytic cracking (FCC) is one of the burning kinetics. The objective of FCC is tomost important processes in the petroleum refinery; maximize the yield of high octane gasoline andemployed for the conversion of straight-run minimize coke formation to make it economicallyatmospheric gas-oil, vacuum residues, and other attractive. 698 | P a g e
  • 2. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714Lee and Grooves [1] developed a mathematical 2. The FCC modelmodel for fluid catalytic cracking using the three- The fluid catalytic cracking (FCC) model equationslump kinetic model of Weekman and Nace [2] for the presented in this study are based on the schematiccracking reactions in the riser-reactor, where (1) gas- flow diagram of the process depicted in Fig. 1. Theoil (the feed and taken as one lump) is converted into feed (gas-oil) enters the riser-reactor from the(2) light gases plus coke lump, and (3) gasoline. The bottom, and is cracked into various products in themajor disadvantage of this model is that it lumped presence of a catalyst. The particle-separator vesselcoke and light hydrocarbon gases together despite the immediately above the riser acts as a disengagingdifferent characteristics of these components. chamber where vapour products and heavyHowever, the three-lump scheme is simple in components are separated from the catalyst usingdescribing the cracking reactions and is able to stripping steam.determine gas-oil conversion and gasoline yieldindependently. It is assumed that the stripping process completely Several workers [3, 4, 5, 6] later used the removes the hydrocarbon gases adsorbed inside thefour-lump kinetic scheme of Lee et al. [7] which catalyst pellets before the spent catalyst is sent to theaccounts for coke formation on the catalyst, to model regenerator.and simulate fluid catalytic cracking process. In thefour-lump scheme, (1) gas-oil is converted into (2) Hydrocarbon Products Flue Gaslight gases (C1-C4), (3) coke, and (4) gasoline. Themajor disadvantage of this model is that it lumped thehydrocarbon light gases and fuel gases (C1-C2)together, making it impossible to predict the yield ofthese gases independently. Particle Dagde et al. [8] applied five-lump kinetic separator Dilute Regionmodel for the cracking of vacuum gas-oil in which(1) gas-oil is converted into (2) liquefied petroleum Dense Regiongas (LPG, i.e. C3-C4), (3) gasoline, (4) coke, and (5)fuel gases (dry gases, i.e. C1-C2). The riser-reactor COwas modelled as a two-phase fluidized-bed reactor Spent CO2assumed to operate isothermally, but the energy catalyst Emulsion Bubble Phase Phasebalance for the riser-reactor and analysis for catalyst- O2regeneration in the regenerator-reactor were notconsidered. Such a five-lump kinetic scheme allowsindependent predictions of LPG and fuel gases which Heatare very significant and desirable in view of the Fresh Gas Oildomestic, commercial, industrial, and laboratory Catalystapplications of these gases [9]. Other workers [10, 11, 12, 13] have alsopresented models for fluid catalytic cracking using Regeneratedten-lump and eleven-lump kinetic schemes for the RISER catalystcracking reactions in the riser-reactor, where the Airvarious lumps are based on the molecular structure of REGENERATORthe components. The feed was lumped into paraffins,naphthenes, and aromatics, in both its heavy and lightfractions. The products were divided into two lumps, Figure 1: Schematic diagram of fluid catalyticwith the first lump comprising products in the cracking unit.gasoline range, while the second lump consists ofcoke and light hydrocarbon gases. The drawbacks of The regenerator operates as a fluidized-bedthese models include mathematical complications in and consists of two regions, namely an upper “dilutethe modelling procedures and lack of experimental region” and a lower “dense region.” The dilute regiondata to validate the models. is the section between the top of the regenerator and In the present paper, models for the riser and the boundary between the two regions, while theregenerator reactors during catalytic cracking of gas- dense region extends from the boundary between theoil are presented. The coke combustion kinetics of two regions to the exit of the regenerator vessel (seeMorley and de Lasa [14] is employed in the Fig. 1). The amount of solids entrained in the diluteregenerator, while a five-lump kinetic scheme is region is usually very small compared to the totaladopted for the catalytic cracking reactions where the amount of catalyst retained in the regenerator vessel.riser-reactor is modelled as a plug-flow reactor Most of the coke on the catalyst pellets is combustedoperating adiabatically. in the dense region, and full combustion of coke to 699 | P a g e
  • 3. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714CO 2 is assumed in this region. Thus, the effect of lumps, with  i as the mass densities of the different lumps, and T is the total mass density of all thethe dilute region on the overall performance of theregenerator is ignored [4].  five lumps; ss is the effectiveness factor; and  is3. The riser-reactor the catalyst deactivation function. Since the overall3.1. The five-lump kinetic scheme cracking rate is affected by the catalyst activity, its Figure 2 shows the five-lump kinetic scheme effect should be incorporated into the above rateused in this study, where (1) gas-oil is converted into(2) gasoline, (3) liquefied petroleum gas (LPG), (4) equations which is represented by  . Thefuel gases, and (5) coke. This kinetic scheme does not deactivation kinetic model of Weekman and Nace [2]account for secondary cracking of products into coke is chosen in this study because of its simplicity andbecause the kinetic rate constants for such secondary popularity in FCC modelling, and is expressed in thecracking reactions are of order of magnitudes smaller formthan the ones for primary cracking reactions. For this   exp K d t c  (6)reason, all secondary cracking reactions are neglectedin this analysis [15, 16, 17]. t where c is the catalyst residence time which is calculated as VGO (1) V tc  R cat (7) K12 Gasoline (2) K13 V where R is the volume of the riser-reactor, cat is  the volumetric flow rate of catalyst into the riser- K23 K d is the catalyst decay coefficient reactor, and K24 LPG (3) which is related to the reaction temperature in the Arrhenius form K34 K d  K do exp(E / RT ) (8) K14 Fuel gas (4) K do where is the catalyst decay pre-exponential factor, E is the activation energy, R is the K15 universal gas constant, and T is the absolute Coke (5) temperature. But Fo  (CTO ) cat Fig. 2: Schematic diagram of five-lump reaction  cat.(9)scheme. which when substituted into eq. (7) noting that The following rate equations areformulated for the various components of the VR  AR LR , yieldsfive-lump kinetic model shown in Fig. 2 A R  cat L R(r1 )  ( K12  K13  K14  K15 ) y12   ss tc  F o  CTO (1) (10) and then into eq. (6), gives(r2 )  [(K 23  K 24 ) y 2  K12 y12 ] ss   K d  cat AR L R (2)   exp (r3 )  ( K 34 y3  K 23 y 2  K13 y12 ) ss  Fo  (CTO )  (11)(3) where CTO is the catalyst-to-gas oil ratio, LR is the(r4 )  ( K 24 y 2  K 34 y3  K14 y12 ) ss height of the riser-reactor, AR is the cross-sectional(4)(r5 )  ( K15 y12 )   ss Fo area of the riser-reactor, is the mass flowrate of(5)  cat K ij gas-oil, and is the mass density of catalyst. Thewhere are the rate constants for the cracking of effectiveness ( ηss ) for the catalytic factor (ri )lumps i to j ; are the reaction rates with cracking of vacuum gas-oil is evaluated as [18] i, with i  1, 2, 3, 4, 5; tanhhi respect to lumps  ss  (12)yi  i / T hi are the mass fractions of the various 700 | P a g e
  • 4. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714 where hi is a modified Thiele modulus given as Coke deposition on the catalyst does not affect the fluid flow. 1 (n  1)  cat  K i Cin1 In each section of the riser-reactor, the hi  (13) catalyst and gas have the same temperature. aext 2 Deff The coke has the same specific heat as the catalyst. where n is the order of reaction (equal to 2 for gas- The riser dynamic is fast enough to justify plug-flow α ext characteristics and a quasi-steady state model. oil cracking); is the catalyst-specific external Instantaneous vaporization occurs at the entrance of Ki the riser-reactor [4]. surface area; are the kinetic rate constants of the Cracking reactions are completed in the Deff riser-reactor. respective lumps; is the effective diffusivity of Ci Applying the conservation principle to a control gas-oil through the catalyst, and is the molar concentration of the various lumps. volume ( AR dL ) of a plug-flow riser-reactor based α ext on the above assumptions, where dL is the It is important to note that in eq. (13) can be differential height of the reactor‟s control volume, defined using the characteristic dimension, LZ , of gives the mass and energy balances as the Zeolite crystallite (i.e. catalyst particle size) since dy dy  1  U 1  (r1 ) R all the cracking reactions take place in the Zeolite dt dL (17) crystallite with little influence from the matrix of the  ( K12  K13  K14  K15 ) y1  R ss 2 catalyst [19]. Therefore, approximating the crystallite dy2 dy geometry to be a sphere, gives the catalyst-specific   U 2  (r2 ) R external surface area as [17] dt dL (18) 6  [(K 23  K 24 ) y2  K12 y12 ] R ss  ext  LZ dy3 dy (14)   U 3  (r3 ) R which when substituted into eq. (13) yields the dt dL (19) following expression for gas-oil:  ( K 34 y3  K 23 y2  K13 y12 ) R ss LZ 3  cat  ( K 12  K 13  K 14  K 15  K 16 ) C o dy4 dyho    U 4  (r4 ) R 6 2 D eff dt dL (20) (15)  ( K 24 y2  K 34 y3  K14 y12 ) R ss where the subscript “o” indicates gas-oil. It is dy5 dy   U 5  ( r5 ) R  ( K 15 y12 ) R ss (21) expected that diffusion of gas-oil takes place in a dt dL USY Zeolite, and the effective diffusion coefficient dTR   R  o AR D   of gas-oil ( eff ) is given by the Erying equation dL ( Fcat C Pcat  Fo C po ) [20].  2  K12 (H 12 )  K13 (H 13 )    Deff  D p exp E D / RT   y1     K14 (H 14 )  K15 (H 15 )   (16) Dp  y K (H )  K (H ) (22) where is the pre-exponential factor for diffusion,  2 23 23 24 24  ED is the activation energy for diffusion.  y3 K 34 (H 34 )  and       Model of the riser-reactor In the derivation of the mathematical model where  R is the average void fraction in the riser- of the riser-reactor, the following assumptions are TR is the temperature of the riser-reactor, made: reactor, Axial dispersion in the riser-reactor is negligible. U is the riser superficial velocity,Fcat is the mass Catalyst particles have a uniform size in a given differential element, and both gas-oil and gasoline C pcat C po flow rate of catalyst, and are the have identical activity decay function  [21] specific heat capacities of catalyst and gas-oil The riser wall is adiabatic. H ij Feed viscosity and heat capacities of all components respectively, are the heat of reaction for the cracking of component i to j , and L is the variable are constant. Adsorption and dispersion inside the catalysts height of the riser-reactor. particles are negligible. Since gas-oil is the feed which is cracked into the Pressure changes throughout the riser-height are due various products, the mass fraction of gas-oil at the to static head of catalyst in the riser. 701 | P a g e
  • 5. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714 inlet ( L  0 ) of the riser-reactor is unity, while the using a fourth-order Runge-Kutta algorithm using the mass fractions of the products at the inlet are equal to boundary conditions (33). zero because no product is formed at the inlet of the riser-reactor. Also, at the inlet of the riser-reactor, the 4. The regenerator feed temperature is taken to be a reference 4.1. Regenerator combustion kinetics Usually, coke is a mixture of different T temperature ( ref ) equal to 800K [22]. These components (carbon, hydrogen, nitrogen, sulphur, boundary conditions at the inlet of the riser-reactor etc.), but mainly carbon [4] Thus, during catalyst are defined mathematically as regeneration in FCC unit, coke is burnt to produce  y1  1 carbon monoxide and carbon dioxide [11, 13]. Also,  the homogeneous CO combustion reaction taking L  0 :  yi  0 i = 2, 3, 4, 5 (23) T  T place in the bubble-phase is assumed to be negligible  R ref compared with the catalytic CO combustion in the We define the following dimensionless variables emulsion phase [23,4]. The following irreversible L coke combustion reactions occur in the emulsion Z (24) phase of the regenerator [24]. LR TR C  O2  CO 2  Heat KC (34) R  (25) Tref C  2 O2  CO  Heat 1 K C (35) where Z is the dimensionless height of the riser- CO  2 O2  CO 2  Heat 1  K CO (36) θ reactor, and R is the dimensionless riser- KC temperature. Using eqs. (24) and (25) in eqs. (17) – where is the reaction rate constant for coke (22), and noting that K CO burning, and is the reaction rate constant for  Fo U  o = (26) AR  o AR the catalytic CO combustion. Equation (36) is the gives “after-burning” reaction which takes place in the dense region if sufficient oxygen is supplied to dy 1 AR  R LR  o  ( K12  K13  K14  K15 ) y12 ss  0 (27) support it. The reaction which goes to completion in dZ Fo the dense region of the regenerator to fully regenerate dy2 AR  R LR  o the catalyst is called the “controlled-after-burning”  [(K 23  K 24 ) y 2  K12 y12 ] ss  0 (28) dZ Fo reaction. However, CO burning is usually initiated dy3 AR  R LR  o by using a promoter, which is a catalyst that speeds  ( K 34 y3  K 23 y 2  K13 y12 ) ss  0 (29) up the reaction of carbon monoxide to carbon dZ Fo dioxide. The promoter, usually a metal like platinum,dy4 AR  R LR  o is attached to the FCC catalyst during manufacturing.  ( K 24 y 2  K 34 y3  K14 y12 ) ss  0 (30)dZ Fo Platinum-based combustion promoters have been utilized in FCC units to catalyze the oxidation of CO dy5 AR R LR  o  ( K15 y1 ) ss  0 2 (31) CO 2 for over 30 years [25, 26, 27]. The better dZ Fo to the dispersion of platinum, the more effective is the d R  AR  R LR  o  combustion of coke. The rate expressions for the dZ ( FcatC pcat  Fo C po ) Tref component gases in the emulsion-phase are obtained as follows.  2  K12 (H 12 )  K13 (H 13 )     y1   Coke (C): (rC )  K C CCS CO2 (37)   K14 (H 14 )  K15 (H 15 )    y K (H )  K (H )  0 (rO )  K CCS CO 2  K COCCOCO/ 2 1 (32) Oxygen (O2): 2 (38) 2  2 23 23 24 24   y 3 K 34 (H 34 )  Carbon dioxide (CO2):   (rCO )   K CCS CO 2  K COCCOCO/ 2 1     2 (39)2 and the boundary conditions (23) become Carbon monoxide (CO):  y1  1 (rCO )  K CCS CO 2  K COCCS CO/ 2 1 2 (40) Z  0 :  y 2  y3  y 4  y5  0  (33) (r ) C is the rate of coke combustion; i (r )   1 where  R are the rates of combustion of individual components  i  O 2 , CO, where o is the volumetric flow rate of gas- of the flue gases (i.e. exit gases), with oil. Equation (27)–(32) were solved numerically 702 | P a g e
  • 6. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714CO 2 ; CCS is the molar concentration of coke on The homogeneous combustion reaction taking place in the bubble-phase is negligible Ci compared with the catalytic CO combustion in thespent catalyst; are the molar concentrations of the   flue gases; K , K , and K , are the rate constants emulsion-phase [14, 4]in eqs. (37) – (40) defined as follows [4, 5] Mass balance for coke combustion in the emulsion-  1  phase K   K C   1 for CO balance (41) Application of the law of conservation of mass to coke combustion in the emulsion-phase based on the   K    K C above assumptions, gives   1  R y R  S y S   rC  GSS  S for CO2 balance (42) dyS L y dLGI    2  d LGI AG (1   E )U a Ua LGI dK    K C (47)  2  2  for O2 balance (43) where the dimensionless rate of change of catalyst- bed height is obtained as where   CO2 / CO is the intrinsic ratio dLGI S   Rof carbon dioxide to carbon monoxide. The mole  d AG  cat (1   E )U a (48)fractions of the various flue gases are expressed withrespect to the molar concentration of oxygen in the The dimensionless variables in eqs. (47) and (48),feed air, in the form and the volume of the regenerator, are defined as CO S R  yO  yS  yR  ; VG  AG LG 1   E ; 2 2 CO  cat ;  cat 2f (44) LG U t  y CO  C CO LGI  ;  a C O2 f LGSS LGSS (49) (45) C CO yS yR  y CO  2 where is the mass fraction of spent catalyst, CO 2 (46) 2f is the mass fraction of regenerated catalyst, S is where CO2 f is the molar concentration of the mass density of spent catalyst, R is the massoxygen in the feed air, and the prime indicates mole  density of regenerated catalyst, cat is the massfraction. VG density of catalyst, is the volume of the4.2. Model of the regenerator-reactor Here, we are concerned with only the dense regenerator,  is the dimensionless time, LGI is theregion since the effect of the dilute region on the LGdynamics of the regenerator is ignored as indicated dimensionless catalyst-bed height, is catalyst-bedabove. The dense region is divided into a bubble- AGphase and an emulsion-phase [28]. The emulsion- height, is the cross-sectional area of thephase is assumed to be a bed at minimum fluidization regenerator, E is the void fraction in the emulsion-velocity, and coke combustion reactions occur in thisphase. The air distributors, spent catalyst, and LGSS phase, is the steady-state catalyst-bed height,cyclones recycles pipes in the emulsion-phase   and S and R are the volumetric flowrates ofproduce enough turbulence which justifies this phaseto be modelled as a continuous stirred tank reactor. spent and regenerated catalyst respectively.The bubble-phase, on the other hand, is dominated bygases at high velocity compared with the velocity in Mass balance for gases in the bubble-phasethe emulsion-phase. Hence, the bubble-phase moves In the bubble-phase, steady-state operationas plug-flow and exchanges mass and heat with the is assumed because of the high velocity of gases, andemulsion-phase without coke combustion reaction it is also assumed that no combustion reaction takesdue to its deficiency in catalyst particles [6]. In the place in this phase due to low density of catalyst.derivation of the mathematical model for catalyst Application of the law of conservation of mass toregeneration, the following assumptions are made: gases in the bubble-phase, with the assumptions of no Unsteady state conditions for the energy and accumulation of gases and without coke combustioncoke combustion balances in the emulsion-phase due reactions, gives the material balance for the flueto the high density of catalyst [4], and Steady state gases ascondition for the same operations in the bubble-phase K L    be GSS 1   B  y ib  y ie  dyibdue to low solid density.  U  dLGI  a  (50) 703 | P a g e
  • 7. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714 K be Application of the law of conservation of where is the mass-transfer coefficient energy to coke on spent and regenerated catalysts as between the bubble and emulsion phases, B is the well as the gases in the emulsion-phase, with coke combustion reactions, accumulation of gases, and yie yib transfer of heat between the bubble and emulsion void fraction in the bubble-phase, and are mass fractions of gases in the emulsion and bubble phases, gives the energy balance in this phase as phases respectively and are defined as dTe (ieo  ieo C p ieo Tao  ie  ie C pieTe)     d U a AG LGI (1   E ) (  ie C pie   R C p ) yib  ib , y ie  ie R  TG  TG (51) ( S C p  S TR   R C p  R Te)   S R with i  O 2 , CO , CO 2 ; ib and ie are the U a AG LGI (  ie C pie   R C p ) R mass densities of gases in the bubble and emulsion n   (H ie )(ri )  O  (H C )(rC )  R LGSS  TG 2 i 1  phases respectively, and is the total mass (  ie C pie   R C p U a Tref ) R density of gases in the regenerator. H be LGSS (Tb  Te) T  dLGI   e (55) Mass balance for gases in the emulsion- (  ie C pie   R C p ) U a LGI d R phase with Application of the law of conservation of mass to Tao T gases in the emulsion-phase, with accumulation of  Tao   , TR  R (56) Tref Tref gases and coke combustion reactions, gives the material balance for the flue gases as   where ieo and ie are the volumetric flowdyie U ieo ( yieo  yie )   ri  GSS K be  yib  yie  GSS (52) L L rates of flue gases at the inlet and outlet of the d LGI (1   E ) U a Ua Ua  Tao emulsion-phase; is the dimensionless inlet air U ieo   where is the incipient velocity in the temperature of the regenerator, ieo and ie are the y mass densities of the flue gases at the inlet and outlet emulsion-phase, ieo are the initial mass fractions of the respective gases in the emulsion-phase, and  O2 of the emulsion-phase respectively, is the mass (ri ) are the reaction rates of the flue gases in the C pieo emulsion-phase. density of gas (oxygen) in the emulsion-phase, C Energy balance in the bubble-phase and pie are the specific heat capacities of the flue Application of the law of conservation of gases at the inlet and outlet of the emulsion-phase energy to gases in the bubble-phase based on the C pS C pR above assumptions, gives the energy balance in this respectively, and are the specific heat phase as capacities of spent catalyst and regenerated catalyst dTb H (T   Te) (1   B ) LGSS ie are ΔH  be b respectively which are taken to be equal, dLGI Ub the heat of reaction of the flue gases in the emulsion- (53) with ΔH C phase, is the heat of reaction for coke combustion,  T T Tb  b , Te  e is the specific area for heat transfer Tref Tref (54) between the bubble and emulsion phases, and R is T Tb Te the dimensionless temperature of the riser. Note that where, and are the dimensionless temperatures in the bubble and emulsion phases the inlet temperature of the regenerator is partly the Tb Te temperature of the spent catalysts entering the respectively, is the bubble-phase temperature, regenerator from the riser, and partly the inlet-air Tref temperature. is the temperature in the emulsion-phase, is a H 5. Hydrodynamic specifications reference temperature taken to be 960K [22], be is the heat-transfer coefficient between the bubble and The interchange mass-transfer coefficients between the bubble and emulsion phases are related Ub in the form [28] emulsion phases, and is the bubble velocity. 1 1 1   Energy balance in the emulsion-phase K be K ce K bc (57) with 704 | P a g e
  • 8. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714   mf DU br  1/ 2 yO   yO  1   yO K ce  6.77  (65) 2 2b 2e  3  y CO   y CO  1   y CO  db  (58) 2 2b 2e (66) and yCO   yCOb  1   yCOe 4.5U mf 5.85D 1 / 2 g 1 / 4 (67) K bc  db  db / 4 5 Te   Tb  1   Te (68) (59) where K ce U mf where is the mass-transfer coefficient between   1 K bc Ua (69) the cloud and emulsion phases, is mass-transfer db and the subscripts e and b indicate emulsion-phase coefficient between the bubble and cloud phases, and bubble-phase respectively. is the effective bubble diameter, D is air diffusivity through the catalyst, g is the acceleration due to 6. Estimation of kinetic parameters 6.1. Kinetic parameters for cracking reactions  gravity, mf is the voidage at minimum fluidization, in the riser-reactor U The riser-reactor model equations contain and br is the rise-velocity of a single bubble in the unknown kinetic parameters such as the reaction rate bed and is given by [29] Ki Ubr  0.711( gdb )1/ 2 constants ( ) for the various reaction paths, and the catalyst deactivation function (  ). These constants (60) Accordingly, the interchange heat-transfer coefficients between the bubble and emulsion phases have to be determined before eqs. (27)-(32) can be may be expressed as integrated. The kinetic parameters for the cracking 1 1 1 reactions based on the five-lump kinetic model are   presented in Table 1. H be H ce H bc (61) with (Kuni and Levenspiel, 1991) The reaction rate constants are functions of 4.5U mf  Go C pg 5.85( K Go  Go C pg )1 / 2 g 1 / 4 temperature and are generally given by the ArrheniusH bc   relation [18] db db / 4 5 (62)   Ei   C K i  K io exp   where Go is the density of the gas mixture, pg  RT  (70) K Go is the specific heat capacity of the gas mixture, Ki where the prime indicates values of predicted by H is the thermal conductivity of the gas mixture, ce K io is the heat-transfer coefficient between the cloud and eq. (70), are the preexponential kinetic H E emulsion phases, and bc is the heat-transfer constants for the respective lumps, and i are the coefficient between the bubble and cloud phases. By activation energies of the different lumps. The rate K bc H bc Ki comparing the expressions for and above, constants ( ) and the stoichiometric coefficients H ce V a corresponding expression for may be ( ij ) of the various lumps are expressed as [30] K M obtained through the expression for ce as  K 1  Vog K 1 ; Vog  o (71) Mg  K Go  Go C pg mf U br  1/ 2  H ce  6.77  3   K 2  VoLPG K 2 ; VoLPG  Mo  db  (63) M LPG (72) The flow velocities in the analysis are related in the Mo form (Froment and Bischoff, 1990)  K 3  Vod K 3 ; Vod  (73) Md U b  U a  U mf  U br (64) Mo  K 4  Voc K 4 ; Voc  (74) U mf Mc where is the minimum fluidization velocity in the emulsion-phase of the regenerator, and Mg  K 5  VgLPG K 5 ; V gLPG  (75) Ua M LPG is the superficial inlet-air velocity into the regenerator. The exit concentrations of the flue gases Mg and exit temperatures from the emulsion and bubble  K 6  Vgd K 6 ; V gd  (76) Md phases are given as [30] 705 | P a g e
  • 9. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714 M LPG 6.2. Kinetic parameters for catalyst  K 7  VLPGd K 7 ; VLPGd  (77) regeneration Md Equation (70) remains valid for estimating Mo Mg the reaction rate constants, K c and K co , used in the where is the molecular weight of gas-oil, is M LPG coke combustion reactions in the regenerator. Table 3 the molecular weight of gasoline, is the shows the kinetic parameters for coke and carbon Mc monoxide. molecular weight of LPG, is the molecular Md Table 3. Kinetic parameters for coke burning weight of coke, and is the molecular weight of dry gas. [14] Table1: Estimated kinetic parameters for cracking REACTION P RE - A C T I VAT I O N reactions in the riser-reactor [30] E X PO NE NT I A L E NE R G YREAC ACTIV. PRE- HEAT STOICHI C O N ST AN T ( k J/ k m o l )TION ENERGY EXPON. OF OME.COE Coke 1.4  10 8 (m 3 / 224.99PATH (kJ/kmol) FACTOR REACT. FFICIENT combustion kmol s) (s-1) (kJ/kg) (Vij) CO catalytic 247.75 70.74Gas-oil 0.02 46.24 -60780 3.2767 Combustion (m 3 ) 1.5 kmol 0.5 /kgto sgasoline 7. Reactors dimensions, feedstock andGas-oil 0.00184 59.75 -2000 8.2655 catalyst propertiesto The dimensions of the riser and regeneratorLPG reactors, as well as properties of the feedstock andGasolin 0.00184 46.24 149000 20.7527 products of the FCC process, were obtained from thee to New Port Harcourt Refinery Company [22] and aredry gas presented in Table 4 – 8.Gasolin 0.00581 59.75 107600 20.965e to Table 4. Dimensions of industrial riser andcoke regenerator reactors [22]Gasolin 0.005566 78.49 200 2.5225e to REACTOR HEIGHT DIAMETERLPG (m) (m)Gasolin 0.002183 78.49 200 6.4022 Riser 22.9 2.9e todry gas Regenerator 35.45 9.8LPG to 0.03174 59.75 100 2.5380dry gas Table 5. FCC feed and products properties [22] API SPECI COMP FLOW-Catalys 83806.556 117705 COMPO GRAVI FIC OSITI RATEt NENT TY GRAVI ON (kg/hr)decay TY (wt. %) Gas oil 21.2 0.927 100 244090 Table 2 shows the average molecular weights of the feed five lumps used in the study. Fuel Gas - - 5.4 13180 Table 2: Average molecular weights of five-lump C3 - - 6.3 15388 kinetic scheme [*31; ** 32]. (LPG) C4 - - 10.7 26118 LUMP AVERAGE MOLEC. (LPG) WEIGHT(kg/kmol) Gasoline 60.0 0.739 45.9 112037 Light 14.0 0.973 17.8 43448 Gas–oil M o  386 * cycle oil Gasoline M g  117 .8 * * Bottoms 0.5 1.072 8.8 21480 Coke - - 5.1 12448 LPG M LPG  46 .7 * * Dry gas M d  18 .4 * * Coke M c  400 * * 706 | P a g e
  • 10. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714Table 6. Physical properties of gas–oil [22] Table 10. Comparison between model-predictions PARAMETER VALUE and plant data for the regenerator Vapour density (kg/m3) 9.52 PARAMETER MODEL PLANT Liquid density at 288oK (kg/m3) 924.8 PREDICTION DATA Specific heat (gas), (kJ/kg K) 3.3 Temperature (K) 1015.07 1016.48 Specific heat (liquid), (kJ/kg K) 2.67 Coke (wt. %) 0.00686 0.007 Heat of vapourization (kJ/kg) 156 O2 (mol. %) 0.0312 0.03 Vapourization temperature (K) 698 CO2 (mol. %) 0.161 0.16 Flowrate (kg/s) 68.05 CO (mol. %) 0.0568 0.00 Inlet temperature (K) 797 It may be seen from Tables 9 and 10 that the model-Table 7. Physcial properties of catalyst [33, 22] predictions compare very well with the plant data, PARAMETER VALUE indicating that the models presented for the riser and regenerator reactors are adequate. For proper catalyst Particle size (m) 6 75  10 regeneration with low carbon content on the Specific heat capacity (kJ/kg K) 1.12 regenerated catalyst and complete burning of CO to Mass flowrate from riser to 1729750 CO2, there must be excess oxygen concentration of 1 regenerator (kg/hr) to 4 mol.% [34] in the regenerator which is consistent Bulk density (kg/m 3 ) 975 with the plant-value of oxygen and model-prediction Mass flowrate (kg/s) 480.49 in Table 10. Although carbon monoxide was not Inlet temperature (K) 975 detected in the flue gases of the plant data, the model- Holdup in the regenerator 5000–70000 predicted 5.7 mol.% of carbon monoxide in the flue (kg) gases is recommended for combustion in the CO boiler to generate superheated steam and energy forTable 8. Physical properties of air [4]. the plant. PARAMETER VALUE Figure 3 shows the variations of the mass fractions of gas-oil and products of the cracking Density (kg/m 3 ) 1.03 process along the height of the riser-reactor. The Specific heat capacity (kJ/kg K) 1.206 mass fraction of gasoline increases with the Bubble-emulsion 0.5 height of the riser-reactor to a maximum value of mass-transfer coefficient (s–1 ) about 0.45 corresponding to dimensionless riser- Bubble-emulsion 0.84 height of about 0.55, where the total heat-transfer coefficient (kJ/m 2 s dimensionless height of the riser-reactor is unity ( K) =1); thereafter, the mass fraction of gasoline Flowrate (m 3 /s) 43.2466 remains approximately constant at the maximum Inlet temperature (K) 370 value as the dimensionless riser-height increases beyond 0.55. The mass fractions of LPG, fuel8. Results and Discussion gases, and coke, increase steadily as the height of Table 9 shows the comparison between riser-reactor increases, while the mass fractionmodel-predictions and plant data for the riser- gas-oil decreases exponentially along the heightreactors, while the comparison between model- of the riser-reactor as it is cracked into thepredictions and plant data for the regenerator-reactor various products.is shown in Table 10.Table 9. Comparison between model-predictions andplant data for the riser- reactor PARAMETER MODEL PLANT PREDICTION DATA Weight fraction of 0.1704 0.17 LPG (C3-C4) Weight fraction of hydrocarbon fuel 0.0546 0.054 gases (C1-C2) Weight fraction of 0.051 0.0511 coke Weight fraction of gas oil 0.2654 0.266 Outlet temperature 652.7 658.00 of riser (K) 707 | P a g e
  • 11. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714 1 0.9 Gas-Oil Gasoline 0.8 LPG Fuel Gases Coke 0.7Mass Fraction 0.6 0.5 0.4 0.3 0.2 0.1 0 0 0.2 0.4 0.6 0.8 1 Dimensionless Height of Riser-Reactor Figure 3. Variations of mass fractions of gas–oil, gasoline, LPG, fuel gases, and coke, along dimensionless height of riser-reactor Figure 4 shows plots of mole fraction of the Figure 4. Variation of mole percent of flue gases with flue gases against dimensionless time, indicating that dimensionless time the concentration of oxygen increases rapidly with time to a maximum value of 0.0742 at a In Fig. 4, the concentration of carbon dimensionless time of 0.025, and then decreases dioxide increases continuously with time, while the continuously to 0.0312 at a dimensionless time of concentration of carbon monoxide increases to a unity. We note in Fig. 4 that initial concentrations of maximum value and then decreases from the oxygen, carbon monoxide, and carbon dioxide were maximum value with time. Although the decrease in chosen [5] in order to simulate the models. Thus, the concentration of carbon monoxide with time from the initial rapid increase in the concentrations of theses maximum value may be due to its conversion to gases in Fig. 4 may indicate that the chosen initial carbon dioxide, this conversion process cannot be concentrations of the gases are less than their „actual‟ attributed to the presence of significant quantity of concentrations at the beginning of coke combustion oxygen in the regenerator since oxygen also As the combustion process progresses, oxygen from decreases with time during the period of carbon the inlet air reacts with coke on the spent catalyst to monoxide depletion. Hence, the decrease in produce carbon monoxide and carbon dioxide, so the concentration of carbon monoxide with time in Fig. 4 concentration of oxygen in the exit gases decreases after attaining the maximum value may be due to the with time while the concentrations of carbon contribution of combustion promoter which sustains monoxide and carbon dioxide increase with time. and speeds up the reaction of carbon monoxide to carbon dioxide. Figure 5 shows the variations of the amount of coke burnt in the regenerator, as well as riser and regenerator temperatures, with dimensionless time. The amount of coke burnt in the regenerator decreases tremendously with time, which is a result of the decrease in the amount of coke on the spent catalyst with time. . 708 | P a g e
  • 12. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714 Coke Burnt 6.00E -02 1100 0.6 Gas Oil Gasoline Riser Reactor Temperature LPG Fuel Gas 1000 5.00E -02 Regenerator Temperature 0.5 Coke 900 Coke Burnt (wt.%) Mass Fraction Temperature (K) 4.00E -02 800 0.4 700 3.00E -02 0.3 600 2.00E -02 500 0.2 400 1.00E -02 0.1 300 0.00E +00 200 0 0 0.2 0.4 0.6 0.8 1 0 5 10 15 Dimensionless Time Catalyst-to-Gas oil RatioFigure 5. Variations of amount of coke burnt, riser Figure 6. Variation of mass fractions of gas-oil andand regenerator temperatures, products with catalyst-to-gas oil ratio with dimensionless timeThe increase in temperature of the regenerator with Accordingly, the mass fraction of gas-oiltime is due to the exothermic coke combustion decreases slightly as the CTO increases to 3 (seereaction taking place in this vessel, while the Fig. 6), thereafter the mass fraction of gas-oildecrease in temperature of the riser-reactor with time remains approximately constant. The slightis due the endothermic cracking reactions. decrease in the mass fraction of gas-oil for CTO  3 is a result of its conversion to8.1 Sensitivity Analysis products. The catalyst spends less time in the A simulation model can be used to optimize riser-reactor at high CTO than at low CTO, whichplant performance by choosing the optimal set of at high CTO reduces the contact time of gas-oiloperating conditions. However, before optimizing and catalyst for effective cracking of gas-oil; thissuch a process, it is important to determine how effect is evident in Fig. 6 where conversion ofsensitive a process is with respect to decision gas-oil and yield of products remain practicallyvariables. In this section, the effects of catalyst-to-gas constant for CTO  3. Thus, an optimum valueoil ratio (CTO) and inlet-air velocity on the of CTO  3 is obtained for the catalyticperformance of the FCC unit are investigated. cracking process. Figure 6 shows the effect of catalyst-to- Figure 7 shows the effect of CTO on thegas oil ratio (CTO) on the conversion of gas-oil riser and regenerator temperatures, indicating that theand yield of products. The mass fraction of all the riser-temperature dropped from an initial value ofproducts increases slightly as the CTO increases about 657.5K to about 653K as the CTO increasesto 3, thereafter the mass fraction of each product from 1 to 3; thereafter, the riser-temperature remainsremains constant for CTO  3. Increasing the uniform for CTO > 3. The initial drop in riser-CTO means increasing the flowrate of catalyst temperature for CTO < 3 is due to increase in theentering the riser-reactor. With more catalyst endothermic cracking of gas-oil within this range ofavailable in the riser, the number of active sites of CTO. The rate of cracking of gas-oil is maximum atcatalyst for the cracking reactions also increases CTO = 3 corresponding to a minimum mass fractionresulting in increased conversion of gas-oil into of gas-oil (see Fig. 6), beyond which, the rate ofproducts for CTO  3. cracking of gas-oil remains constant at the maximum value resulting in a constant temperature of the riser- reactor as obtained in Fig. 7 for CTO > 3. 709 | P a g e
  • 13. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714 about 14m/s, beyond which, the regenerator- 670 1200 temperature decreases as the inlet-air velocity increases. The increase in the amount of coke burnt and the initial increase in regenerator-temperature as 665 1000 the inlet-air velocity increases, are due to increase in Regenerator Temperature (K) Riser Temperature (K) the rate of the exothermic coke combustion reactions Riser Temperature 660 800 resulting from increased rate of supply of oxygen Regenerator Temperature from the inlet-air. At the maximum regenerator- temperature (obtained in Fig. 8 to be about 1080K), 655 600 the operation of the regenerator is said to have reached „total combustion regime‟ and any carbon 650 400 monoxide present in the regenerator is converted to carbon dioxide. For inlet-air velocity above 14m/s, the spent catalyst spends less time (low residence 645 200 time of spent catalyst) in the regenerator caused by channelling and by-passing effect inherent in typical 640 0 fluidized-bed reactors [35] 0 5 10 15 Catalyst-to-Gas oil Ratio Coke Burnt 1150 1.40E -02Figure 7. Variation of riser and regenerator Regenerator Temperature 1100temperatures with catalyst-to-gas oil ratio Regenerator temperature (K) Coke Burnt (wt.%) 1.20E -02 1050 Thus, the optimum riser-temperature for 1.00E -02maximum cracking of gas-oil is about 653K. We notethat the optimum temperature of the riser-reactor can 1000 8.00E -03vary significantly depending on the processconditions. It is interesting to observe that the 950 6.00E -03optimum riser-temperature of about 653K obtained inFig. 7 corresponds roughly to the temperature at the 4.00E -03 900point of intersection of the riser and regeneratortemperatures in Fig. 5. In other words, maximum 2.00E -03 850conversion of gas-oil and maximum yield of gasolineoccur at the optimum riser-temperature (obtained 0.00E +00 800here to be about 653K) which in the present analysis 12 14 16 18 20corresponds to dimensionless riser-height of about0.55 (see Fig. 5), and is consistent with Fig. 3 for Air Velocity (m/s)gasoline yield. Figure 7 also indicates that the Figure 8. Variation of coke burnt and regenerator-regenerator-temperature remains practically constant temperature with air velocity.at about 1000K for all values of CTO, meaning thatCTO does not have significant effect on the Some portion of the spent catalyst alsotemperature of the regenerator. This is because CTO escapes with the flue gases without proper contactmainly influences the catalytic cracking reactions in with the combustion air. All these factorsthe riser-reactor with attendant effect on the riser- associated with high inlet-air velocity result intemperature for CTO < 3. The spent catalyst enters heat losses which reduce the temperature of thethe regenerator at a temperature equal to the outlet regenerator and this effect increases as the inlet-temperature of the riser-reactor of about 653K air velocity increases beyond 14m/s, as obtained(model-estimated value, see Table 9). This in Fig. 8. Even though the amount of coke burnttemperature of spent catalyst is less than the and the rate of the exothermic coke combustionregenerator-temperature of about 1000K; hence, the reactions increase as the inlet-air velocitytemperature of the regenerator is not affected by the increases, the overall quantity of heat generatedtemperature of spent catalyst, and remains constant in the regenerator for inlet-air velocity greaterirrespective of the CTO. then 14m/s may be less than the quantity of heat Figure 8 shows the effect of inlet-air lost from the regenerator to the surrounding;velocity on the regenerator-temperature and the hence, the decrease in regenerator-temperature foramount of coke burnt. As the inlet-air velocity inlet-air velocity greater than 14m/s. It is veryincreases, the amount of coke burnt in the regenerator essential that the coke be burned off the spentincreases continuously, but the regenerator- catalyst at the same rate as it is produced in thetemperature initially increases to a maximum value of riser-reactor. This can be achieved by maintainingabout 1080K corresponding to inlet-air velocity of a small amount of excess oxygen in the 710 | P a g e
  • 14. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714regenerator above that which is required to burn C piethe coke. When all the coke is not burnt, the unit specific heat capacities of flue gases at theis said to be „behind-in-burning‟ with the result outlet of the emulsion-phase, kJ/kg K.that the catalyst turns grey [34] and looses its C pgactivity thereby decreasing the yield of desired specific heat capacity of gas mixture used inproducts in the riser. To avoid this, the velocity eq. (62), kJ/kg Kof air entering the regenerator should be C pSincreased gradually. specific heat capacity of spent catalyst, kJ/kg K.9. Conclusion C pR Models which describe the operations of the specific heat capacity of regeneratedriser and regenerator reactors in an industrial FCC catalyst, kJ/kg K.unit have been presented. The five-lump kinetic CTO catalyst-to-gas oil ratioscheme adopted for the cracking reactions in theriser-reactor, and the steady-state and unsteady-state db effective bubble diameter, m.models developed from the mass and energy balances D air diffusivity through the catalyst in thefor the riser and regenerator reactors, give adequate regenerator used in eq. (59), m2/s.predictions of the feed and products of gas-oilcracking and catalyst-regeneration processes. It is Deff effective diffusivity of gas-oil through theshown that CTO and inlet-air velocity have catalyst in the riser-reactor used in eq. (16), m2/ssignificant effects on the performance of the riser and Dpregenerator reactors respectively. The riser- pre-exponential factor for diffusion used intemperature varies from about 650K to 800K, while eq.(16), m2/sthe regenerator-temperature ranges from about 650K EDto 1080K (see Figs. 5, 7, and 8). The optimum values activation energy for diffusion used in eq.of process variables obtained in the analysis for (16), kJ/kmoleffective operation of FCC are inlet-air velocity of E activation energy used in Eq. (70), kJ/kmol.about 14m/s, riser-temperature of about 653K, and Focatalyst-gas oil ratio of 3. mass flowrate of gas-oil, kg/s FcatNotation mass flowrate of catalyst, kg/s g acceleration due to gravity, m/s2.AG cross-sectional area of the regenerator, m2. hiAR modified Thiele modulus used in eq. (12). cross-sectional area of the riser reactor, m2 hoCCS modified Thiele modulus for gas-oil used in molar concentration of coke on spent eq. (15).catalyst, kmol/m3. H beCi heat-transfer coefficient between the bubble molar concentrations of the flue gases (i.e. and emulsion phases, kJ/m2s Kexit gases, i = O2, CO, CO2), kmol/m3. H bcCO2 f heat-transfer coefficient between the bubble molar concentration of oxygen in the feed and cloud phases,air used in eqs. (44)-(46), kJ/m2s K kmol/m3. H ceCie heat-transfer coefficient between the cloud molar concentrations of flue gases in the and emulsion phases,emulsion-phase, kmol/m3. kJ/m2s KCib KC molar concentrations of flue gases in the reaction rate constant for coke burning,bubble-phase, kmol/m3. m3/kmol s.Cp K CO cat specific heat capacity of the catalyst in the reaction rate constant for the catalytic COriser-reactor, kJ/kg K combustion, m3/kmol s.C po Kd specific heat capacity of gas-oil in the riser- catalysts decay coefficient used in eq. (8),reactor, kJ/kg K s-1C pieo K do specific heat capacities of flue gases at the catalyst decay pre-exponential kinetic factorinlet of the emulsion-phase, kJ/kg K. -1 used in eq. (8), s 711 | P a g e
  • 15. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714K be U mf mass-transfer coefficient between the bubble minimum fluidization velocity, m/s. -1and emulsion phases, s Ua superficial inlet air velocity, m/s.K ce U ieo mass-transfer coefficient between the cloud incipient velocity into the emulsion-phase,and emulsion phases, s-1. m/s.K bc U br mass-transfer coefficient between the bubble rise velocity of a single bubble, m/s.and cloud phases, s-1. Ub bubble velocity, m/s.Ki Vij reaction rate constant used in Eq. (70), i to stoichiometric ratio of componentm3/kmol s.K io component j used in eqs (71)-(77). pre-exponential constant used in Eq. (70),m3/kmol s. VG volume of the regenerator, m3.K Go yie thermal conductivity of gas mixture used in mass fractions of gases in the emulsion-eq. (62), W/m K phase.K , K , K , rate constants used in Eqs. (37) – yib mass fractions of gases in the bubble-phase.(40), m3/kmol s. ySLGI mass fraction of spent catalyst. dimensionless catalyst-bed height.LG yR mass fraction of regenerated catalyst. catalyst-bed height, m. yieoLGSS initial mass fraction of the respective gases steady-state catalyst-bed height, m. in the emulsion-LR height of riser-reactor, m phase.Lz  yO2 characteristic dimension of Zeolite mole fraction of oxygen used in eq. (44).crystallite size used in eq. (14), µm  yCO mole fraction of carbon monoxide used inMi i used in molecular weight of component eq. (45).eqs. (71)-(77), kg/kmol  yCO2(ri ) mole fraction of carbon dioxide used in eq. reaction rates of flue gases in the emulsion- (46).phase, kmol/s. yi(  rC ) mass fraction of component i rate of coke combustion reaction; kmol/s. Z dimensionless height of riser-reactor.R universal gas constant, kJ/kmol K.t time, s Greek Letters  intrinsic ratio of carbon dioxide to carbonT absolute temperature, K. monoxide defined after eq. (43).Tb  ext dimensionless temperature in the bubble- catalyst-specific external surface area usedphase. in eq. (14), m-1Te  dimensionless temperature in the emulsion- a dimensionless parameter defined in eq.phase. (69).Tb temperature in the bubble-phase, K. R average void fraction in the riser-reactorTe temperature in the emulsion-phase, K. E void fraction in the emulsion-phase. Tao B dimensionless inlet air temperature of the void fraction in the bubble-phase.regenerator.  mfTR temperature of riser-reactor, K voidage at minimum fluidization.  oTR dimensionless temperature of the riser- mass density of gas-oil, kg/m3 ireactor. mass density of lump i, kg/m3TR riser-reactor temperature, K. T total mass density of all lumps defined after eq. (6), kg/m3 712 | P a g e
  • 16. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714S References mass density of spent catalyst, kg/m3. [1] E. Lee, F, Grooves, Mathematical Model ofR mass density of regenerated catalyst, kg/m3. the Fluidized Bed Catalytic Cracking Plant, Transaction of the Society of Computer cat mass density of catalyst, kg/m3. Simulation, 1985, 2, 3, 219-236.ib mass densities of gases in the bubble-phase, [2] V. W. Weekman, D. M. Nace, Kinetics of Catalytic Cracking Selectivity in Fixed,kg/m3. Moving and Fluidized-Bed Reactors, AIChEie mass densities of gases in the emulsion- Journal, 1970, 16, 3, 397-404. [3] H. Ali, S. Rohani, Effect of Crackingphase, kg/m3. Reaction Kinetics on the Model Predictions ieo mass densities of flue gases at the inlet of of an Industrial Fluid Catalytic Cracking Unit, Chemical Engineeringthe emulsion-phase, kg/m3. Communications, 1996, 149: 163-184.ie mass densities of flue gases at the outlet of [4] H. Ali, S. Rohani, J. P. Corriou, Modelingthe emulsion-phase, kg/m3. and Control of a Riser Type Fluid Catalytic Cracking (FCC) Unit, Transactions of the O2 Institution of Chemical Engineers, 1997, 75, mass density of gas (oxygen) in theemulsion-phase, kg/m3. 401 – 412.TG total mass density of gases in the [5] I. S. Han, C.B. Chung, Dynamic Modeling and Simulation of a Fluidized Catalyticregenerator, kg/m3. Cracking Process. Part 1: Process Modeling, Go mass density of gas mixture used in eq. (62), Chemical Engineering Science, 2001, 56, 1951 – 1971.kg/m3 [6] D. Krishnaiah, V. Gopikrishna, B. Awang,S volumetric flow rate of spent catalyst, m3/s. S. Rosalam, “Steady State Simulation of a Fluid Catalytic Cracking Unit, Journal ofR volumetric flow rate of regenerated catalyst, Applied Science, 2007, 7, 15, 2137-2145.m3/s. [7] L. S. Lee, V. W. Chen, T. N. Huang, Four-ieo volumetric flow rates of flue gases at the Lump Kinetic Model for Catalytic Cracking Process, Canadian Journal of Chemicalinlet of the emulsion-phase, m3/s. Engineering, 1989, 67, 615 – 623.ie volumetric flow rates of flue gases at the [8] K, K. Dagde, J. G. Akpa, Y. T. Puyate, E. O. Oboho, Five-Lump Kinetic Model for Fluidoutlet of the emulsion-phase, m3/s. Catalytic Cracking of Gas-Oil in a Fluidizedo Bed Reactor, Journal of the Nigerian volumetric flowrate of gas-oil used in eq. Society of Chemical Engineers, 23, 1&2, 1-(27), m3/s 19. ISSN: 0794-6759ΔH ie [9] O. Xu, S. E. Hong–Ye, M. U. Sheng–Jing, heat of reaction of flue gases in theemulsion-phase, kJ/kmol. C. H. U. Jan, 7-Lump Kinetic Model for Residual Oil Catalytic Cracking, Journal ofΔH C Zhejiang University Science A, 2006, 7, 11, heat of reaction for coke combustion,kJ/kmol. 1932–1941.H ij i to j , [10] S. Kumar, A. Chadha, R. Gupta, R. Sharma, heat of reaction for cracking lump A Process Simulator for an Integrated FCCkJ/kmol Regenerator System, Industrial and specific area for heat transfer between the Engineering Chemistry Research, 1995, 34,bubble and emulsion 3737-3748 phases, m2/m3. [11] A. Arbel, Z. Huang, I. H. Rinard, R. Shinnar, A. V. Sapre., Dynamic and Control catalyst deactivation function. of Fluidized Catalytic Crackers. 1.R dimensionless temperature of riser-reactor Modelling of the Current Generation of FCCs. Industrial and Engineering ss Chemistry Research, 1995, 34, 1128-1243. effectiveness factor dimensionless time. [12] I. M. Arandes, M.J. Azkoiti, H. I. de Lasa, Modelling FCC Units Under Steady and Unsteady State Conditions Canadian Journal of Chemical Engineering, 2000, 78, 111-123. 713 | P a g e
  • 17. K. K. Dagde, Y. T. Puyate / International Journal of Engineering Research and Applications (IJERA) ISSN: 2248-9622 www.ijera.com Vol. 2, Issue 5, September- October 2012, pp.698-714 [13] R. M. Rao, R. Rengaswamy, A.K. Suresh, [27] T. Gauthier, J. Bayle, P. Leroy, FCC: K. S.Balaraman Industrial experience with Fluidization Phenomena and Technologies, Object Oriented Modelling, FCC Case Oil and Gas Science and Technology, 2000, Study. Chemical Engineering Research and 55, 2, 187-207. Design,” 2004, 82, 527-552 [28] D. Kunii, O. Levenspiel, Fluidization [14] K. Morley, H. I. de Lasa, On the engineering, 2nd ed., 1991 (Butterworth– Determination of Kinetic Parameters for the Heinemann, London.) Regeneration of Cracking Catalyst, [29] G. F. Froment, K. B. Bischoff, Chemical Canadian Journal of Chemical Engineering, reactor analysis and design, 2nd ed., 1990 1987, 65, 773 – 777. (John Willey, London). [15] J. Ancheyta-Juarez, L. F. Isunza, E. A, [30] K. K. Dagde,, Development of Model Rodriguez, 5 – Lump Kinetic Model for Gas Equations for the Simulation of Fluid Oil Catalytic Cracking, Applied Catalysis A, Catalytic Cracking Reactors PhD Thesis, 1999, 177, 2, 227 – 235. Department of Chemical/Petrochemical [16] L. L. Oliveira, E. C. Biscaia, “Catalytic Engineering, Rivers State University of Cracking Kinetic Models, Parameter Science and Technology, Port Harcourt, Estimation and Model Evaluation Industrial Nigeria, 2009. and Engineering Chemistry Research, 1989, [31] J. Ren, H. X. Weng, F. Y. Liu, Study on 28, 3, 264 – 271. Composition and Structure of Heavy Oil, [17] M. Al-Sabawi, A. A. Jesus, H.I. de Lasa, Petroleum Processing, 1993, 24, 9, 56 – 61. Kinetic Modeling of Catalytic Cracking of [32] F. C. Peixoto, J. L. Medeiros, Reaction in Gas Oil Feedstocks: Reaction and Diffusion Multi-indexed Continuous Mixtures: Phenomena, Industrial and Engineering Catalytic Cracking of Petroleum Fractions, Chemistry Research, 2006, 45, 1583 – 1593. AIChE Journal, 2001, 47, 4, 935 – 947. [18] O. Levenspiel, Chemical Reaction [33] E. O. Oboho, J. G. Akpa, Modelling of a Engineering, 2001, 3rd ed., (John Wiley, Fluid Catalytic (FCC) Riser Reactor: The New York). Four-Lump Model, Journal of Modeling, [19] G. Tonetto, J. A. Atias, H. I. de Lasa, FCC Design and Management of Engineering Catalysts with Different Zeolite Crystallite Systems, 2002, 1, 1, 39–52 Sizes: Activity, Structural Properties and [34] D. Bai, J. X. Zhu, Y. Yin, Z. Yu, Simulation Reactivity, Applied Catalysis A, 2004, 40, 9- of FCC Catalyst Regeneration in a Riser 25 Regenerator, Chemical Engineering [20] D. M. Ruthven, Principles of Adsorption Journal, 1998, 71, 97 – 109. and Adsorption Processes, 1984, (John [35] N. P. Cheremisinoff, P. N. Cheremisinoff, Willey, New York). Hydrodynamics of Gas-Solid Fluidization, [21] J. S. Ahari, A. Farshi, K. Forsat, A 1984, (Gulf Publishing Company, Houston). Mathematical Modeling of the Riser Reactor in Industrial FCC unit, Petroleum and Coal, 2008, 50, 2, 15 – 24. [22] NPHRC, New Port Harcourt Refinery Training Project: Specific for Junior Staff,” Area 3, Process Description, 1, (Comerint, 1987). [23] S. S. E. H. Elnashaie, S. S. Elshishini, Digital Simulation of Industrial Fluid Catalytic Units – IV: Dynamic Behaviour, Chemical Engineering Science, 1993, 48, 567 [24] P. B. Weisz, R. B.Godwin, Combustion of Carbonaceous Deposits within Porous Catalyst Particles, Part II: Intrinsic Burning Rate, Journal of Catalysis, 1966, 6, 227-236 [25] W. C. Yang, Handbook of Fluidization and Fluid Particle Systems, 2003, (CRC Press, New York, ISBN0-8247-0259-X) [26] J. H. Gary, G. E. Handwerk, Petroleum Refining: Technology and Economics, 2001, 4th ed., (CRC Press, London, ISBN 0-8247- 0482-7). 714 | P a g e

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