Hydrodynamics of Riser Units and Their Impact on FCC Operation

4504

Ind. Eng. Chem. Res. 1997, 36, 4504-4515

Hydrodynamics of Riser Units and Their Impact on FCC Operation Ce´ line Derouin,† David Nevicato,† Michel Forissier,‡ Gabriel Wild,§ and Jean-Rene´ Bernard*,† ELF Aquitaine, Centre de Recherche ELF de Solaize, B.P. 22, 69360 Solaize, France, Laboratoire de Ge´ nie des Proce´ de´ s Catalytiques, CPE, B.P. 2077, 69616 Villeurbanne Cedex, France, and Laboratoire des Sciences du Ge´ nie Chimique, CNRS ENSIC, B.P. 451, 54001 Nancy Cedex, France

The hydrodynamic characteristics of the riser of a FCC plant are investigated in a cold pilot plant. The results are checked and completed with measurements made in real operating FCC units. In both cases, tracing of the gas, sampling probes, and γ-tomography were used. The gas and solid velocity profiles are described by using a plug-flow model with radial dispersion. The gas velocity profile is formally represented by an Ostwald-de Waele type equation. The axial solid concentration profile is fitted to data obtained in a commercial FCC unit. The radial dispersion coefficient of the gas is directly measured in the plant. The kinetics of the reactions occurring during catalytic cracking are investigated in a microreactor and are represented by a 19-lump model including deactivation by coking of the catalyst. A combination of this reaction model with the hydrodynamic model is able to predict with a good precision the yields of the different product families obtained in an industrial FCC unit. It is checked by comparison with sampling data in the plant, and it yields useful information on the influence of riser hydrodynamics on yields and products quality. The two main applications of circulating fluidized beds are oil and petrochemical processing, on the one hand, and combustion, on the other hand. Understanding of such reactors is in constant progress; however, information is lacking when large solid fluxes (> 100 kg‚m-2‚s-1) are used, and experimental techniques to investigate the local properties are scarce when the solid concentration is high (Berruti et al., 1995). Prediction of the conversion by catalytic reaction in a riser has been made in only very limited cases (e.g., by Schoenfelder et al., 1994, 1996). The aim of this work is 3-fold: (a) Determine hydrodynamic characteristics relevant for a FCC riser. (b) Combine these data with the results of an investigation of the kinetics of catalytic cracking to determine the yields of the different cracking products. (c) Check the results of the model with data obtained on a commercial riser. The systematic investigation of the hydrodynamics of the riser was done in a cold setup; however, a number of measurements were made in FCC plants of the company ELF ANTAR FRANCE. Table 1 shows characteristics of the cold pilot plant and typical data from an industrial riser. Figure 1 shows two different commercial FCC plants, a side by side and a stacked unit. Experimental Techniques The following measuring techniques were used systematically in the pilot plant but also in some cases in industrial plants: (a) Tracers of the gas and the solid: In the cold plant, a continuous flow of a non adsorbable tracer gas (helium) was injected axially at the bottom of the column, and the radial profile of the tracer gas concentration was measured at different levels of the riser. Details on these measurements may be found in Martin †

Centre de Recherche ELF de Solaize. CPE. § CNRS ENSIC. ‡

S0888-5885(97)00432-6 CCC: $14.00

et al. (1992a). Pulses of radioactive tracers of gas and catalyst were also injected at the riser bottom to estimate their velocity and axial dispersions (Bernard et al., 1989). (b) Suction probes according to van Breugel et al. (1970): By sampling of the gas-solid flow (both isokinetic and nonisokinetic), it was possible to determine simultaneously the local gas and solid mass flow rates at different axial and radial locations in the riser. The results thus obtained are consistent with the total volume flow rate of the gas, as shown in Figure 4. However, discrepancies appear for the solid flux at high values of the solid flux; this may be due to choking in the sampling probe. In the industrial plant the local axial velocity of the solid was determined by using Pitot probes (Azzi et al., 1990). (c) γ-Tomography: This technique allows one to determine the axial and radial solid concentration profiles in the cold plant as well as in the industrial riser. Details are found in Martin et al. (1992a) and in Turlier and Bernard (1992). (d) Sampling reactive mixtures of catalysts and cracked hydrocarbons was done on commercial risers at various levels and various radial positions at each given level. This mixture was quickly quenched in the receiver: gas residence time in the sampling line was less than 0.1 s at riser temperature, so that further cracking was estimated to be negligible in the sampling setup. A permanent flow of helium was also injected in the center line of a commercial riser, and gases were sampled 15 m downstream at different radial positions. Hydrocarbons and helium analysis could provide quantitative information on gases’ radial dispersion and composition gradients. Description of the Hydrodynamics in the Riser The description of the hydrodynamics is made with a type II model, according to the classification proposed by Harris and Davidson (1994) and by Berutti et al. (1995). These models consider generally a two-zone description with a central core in which gas is flowing © 1997 American Chemical Society

Ind. Eng. Chem. Res., Vol. 36, No. 11, 1997 4505

Figure 1. Side by side and stacked FCC plants. Table 1. Characteristics of the Cold Plant and an Industrial Riser cold plant i.d. (m) height (m) solid mean diameter (µm) density (kg/m3) mass flow rate (kg‚m-2‚s-1) gas pressure (MPa) density (kg/m3) mean velocity (m/s)

0.184 11.4 FCC catalyst (Geldart A) 80-300 air atmospheric pressure 1.2 up to 6.7

at high velocity, entraining dilute solid with a small slip velocity generally equal to the particle terminal velocity (Berruti and Kalogerakis, 1989; Ambler et al., 1990; Rhodes, 1990; Kagawa et al., 1990). The second zone is a peripheral annulus where concentrated solid is flowing down, with a gas velocity close to zero. These models are convenient to describe in a simple manner the hydrodynamic characteristics of the bed. However, they are less adapted to describe CFB in which chemical reactions are occurring. The main problem is that stagnant gas is in contact with concentrated solid in the annulus. This is not in agreement with physical phenomena, and a mass-exchange correlation between the core and annulus must be introduced (Rhodes, 1990; Pugsley et al., 1992; Harris and Davidson, 1994). This correlation is extremely difficult to establish. The core-annulus representation is justified by an abrupt solid density change when moving from the wall to the center. But the other variations of the main physical parameters are smooth, as will be shown later, and a unique mass-transfer law can well represent the radial mixing of the gas, as shown by Martin et al. (1992a). Thus, it seems more realistic to adopt a continuous description of the riser reactor in the two dimensions r and z, even if the radial density profile of the solid is described by an abrupt correlation like the one proposed by Zhang et al. (1991). This simple description is also in agreement with the experimentally observed lack of catalyst downflow at the wall of commercial risers (Azzi et al., 1990). Gas Velocity Profiles. The local velocity of the gas is supposed to be vertical with a nonuniform velocity; mixing of the gas is represented by a radial dispersion coefficient Dr; the radial gas velocity profile can formally

FCC riser 0.7-1 ∼30 60-70 1500-1550 300-600 complex mixture of hydrocarbons 0.28-0.33 ∼21 (bottom) ∼5 (top) 4-15

be represented by an Ostwald-de Waele type equation:

Ug(r) ) Ug,av

(3nn ++11)(1 - (Rr ) ) n+1/n

(1)

Figure 2 shows the types of velocity profiles represented by this equation as a function of n. The validity of eq 1 was checked by isokinetic sampling in the cold riser (Figure 3). Experimental gas velocity profiles fit very well with the Ostwald-de Waele correlation; it shows that no downward gas flux can be detected even when solid catalyst is refluxing at the wall. Figure 4 confirms that the technique of gas sampling along the riser cross-sectional area is valid since cylindrical integration of the data (fitted either with a polynomial expression or with the Ostwald-de Waele expression) is in agreement with the superficial velocity. The concavity of the gas velocity profiles increases when the solid mass flux increases and/or when the gas velocity decreases (Figures 3 and 6). Figure 5 shows that the Ostwald-de Waele index n can be extrapolated to 2 for commercial riser conditions. Gas velocity profiles were measured by isokinetic sampling at different elevations of the cold model (halfway up the riser and 2 m upstream of the elbow). It is known that an abrupt elbow provokes significant slackening and concentration increase of the solid, as detailed therein (Martin et al., 1992b). Figure 6 shows that the gas velocity profile is not significantly affected by this concentration increase. No dissymmetry provoked by the latter is detected, although it should appear if measurements were done closer to the top, since the elbow is dissymmetric by nature.

4506 Ind. Eng. Chem. Res., Vol. 36, No. 11, 1997

Figure 2. Radial velocity profiles according to Ostwald-de Waele’s law (eq 1) as a function of the consistency index n.

not be solved because of a helium excess compared to hydrocarbons. This is obviously because a part of the latter remains adsorbed on the catalyst, so that the average concentration of the helium in the gas phase is no longer well-known: Schuurmans (1979) states this amounts ca. 10% by weight, but Martignoni and de Lasa’s models (1995) evaluate it to ca. 50%. Equations 1 and 2 (helium/hydrocarbons balance) find a solution for an adsorption between 10 and 30% by weight on the catalyst, including 6% coke. The lower bound determines n ) 0 (flat velocity profile), while the upper bound determines n ) infinity (sharp triangular profile). Gas Radial Dispersion. The former experiments allow one to conclude that gases do not radially mix, given the very small gas residence time on the 12 m distance between tracer injection and sampling levels. The gas tracer concentration profiles were interpreted by means of a plug-flow model with radial dispersion of the gas. The balance equation of helium can be written as Figure 3. Radial velocity gas profile in the cold plant.

Helium tracing was also used to estimate directly the consistency index n of eq 1. Details may be found in Martin (1990) and Martin et al. (1992b). Knowing the mixing cup concentration Cav of the continuously injected tracer (helium) and assuming the velocity profile to follow Ostwald-de Waele’s law, the mixing cup concentration can also be evaluated by the following equation:

CavQtot ) CavπR2Ug,av )

∫0R2πrUg(r) C dr

(2)

Using the experimental values C(r) of the tracer concentration to evaluate the integral on the right-hand side of eq 2, the consistency index n can be estimated. Helium tracing was also done in an industrial plant. In this case only gaseous hydrocarbons were sampled with a line equipped with a filter, excluding catalyst sampling. These data were processed from the tracer and hydrocarbons material balance. With cold flow experiments (helium in air), this technique worked very well. However, with the experiments made in the industrial riser, the equation could

∂(CUg) Dr ∂ ∂C r )0 ∂z r ∂r ∂r

( )

(3)

Using the consistency index determined either by the helium balance (eq 2) or by actual gas velocity measurements, a fitting of the concentration profiles allows one to estimate the radial dispersion coefficient. The values of the radial dispersion coefficient Dr of the gas were on the order of magnitude 10-3-10-2 m2/s in the cold plant, a value slightly superior to the value obtained without solid circulation perhaps because of increased radial transport by the porous catalyst. Despite the problems mentioned above, the radial dispersion model used in Martin et al. (1992a) was also used to estimate gas dispersion in the industrial plant. A value of 0.03 m2/s was found for n ) 2. This result did not prove to be very sensitive to the n index since its variation is 15% when n varies between 0 and 4. Figure 7 shows the results of the radial dispersion model compared to the experimental data. Catalyst Flux Profiles. The solid flux Fs profiles are not always perfectly symmetric, even far from the entrance or the outlet of the column and with a perfectly vertical column. However, in order to use a two-


Figure 4. Comparison of superficial gas velocity from cold setup flow rate and from gas sampling..

Figure 5. Estimation of the consistency index at the bottom of an industrial riser (Ug ∼ 4 m/s; Fs ∼ 350 kg‚m-2‚s-1).

dimensional model, the conditions of symmetry are assumed for the solid flux using a relationship similar to the one adopted for the gas phase:

Fs(r,z) ) (Fs,av(z) - Fs,min)

(3nn ++11)(1 - (Rr ) ) + n+1/n

Fs,min (4) Figure 8 shows typical radial flux profiles obtained

in the cold setup. In this case, Fs,min is negative since refluxing catalyst is detected at the wall. In fact, in the industrial riser, downflow in the annulus region near the wall was not observed with the Pitot probe (Azzi et al., 1991), so that Fs,min is zero or slightly positive. It represents a parabolic profile shape which could be detected on a commercial plant at high flux (350 kg‚m-2‚s-1). Other shapes with positive humps at relative radius 0.3 were recently shown for


Figure 6. Gas velocity profiles at half riser height and 2 m upstream of the elbow in the cold setup (Ug ) 5.2 m/s).

Figure 7. Results of helium tracing in an industrial riser: model and experimental data.

large solid fluxes in a Workshop on Circulating Fluidized Beds, Tours (France, 1995). This kind of profile was also found on an other commercial riser at 1100 kg‚m-2‚s-1 catalyst flux. The reason for these differences is not clear; they may be attributed to different oil or catalyst distributions. It must, however, be pointed out that the catalyst flux distribution is not a key point to while modeling FCC risers. In effect, flux distribution affects time on stream distribution if dispersion is small, but time on stream

is not a significant factor for catalyst activity. We found that coke is made extremely rapidly and evenly on the catalyst at the riser’s bottom and that catalyst activity depends only on coke content, as discussed therein. Since the latter is constant, only the catalyst hold-up distribution is important from a modeling point of view. However, a flux profile must be determined in order to calculate catalyst concentrations from gas velocities which depend on conversion and from a slip velocity correlation.


Figure 8. Radial mass flux profiles in the cold setup.

In the present work, we consider a regular distribution of Fs which can be represented by eq 4. Catalyst Concentration Profiles. The gammametric measurements give access to the local values and show the classical annulus-core structure: near the wall, the solid concentration is much higher than in the center of the riser. They confirm the observations made earlier by Martin et al. (1992b), Senior and Brereton (1992), and Wong et al. (1992): the mean solid concentration decreases from the bottom to the middle of the height of the riser; at the upper part of the column, 2 m upstream of the abrupt elbow, an increase of the solid concentration is observed, through a reinforcement of the core-annulus structure; however, there was not any discernible effect of the elbow orientation. Figure 9 shows the radial density profiles obtained in the cold setup for Fs ) 303 kg‚m-2‚s-1 and Ug ) 6.7 m/s. They can be represented by an equation similar to Zhang’s (1991) correlation, where is the porosity (gas hold up): b

(r) ) ava(r/R) +c

with Cs,exp ) Fp(1 - )

(5)

Slip Velocity between Gas and Catalyst. The former data confirm and allow one to quantify the

presence of core-annulus structure in the cold model. These data cannot be used directly in a model because the catalyst concentration is the result of hydrodynamic phenomena involving primarily the solid flux and the gas velocity. The latter is an unknown in the model since it depends on volumetric expansion due to cracking. This is why it is necessary to develop a relationship between gas and catalyst velocity. It is presently difficult to use a simple hydrodynamic model established on strong physical bases to estimate the slip velocity, because of the elbow effect. Typical slip velocities on a riser diameter obtained on the cold setup are shown in Figure 10. They are represented as slip velocities Ug - Vs. A slip coefficient is also defined according to

Fs(r,z) ) SL(r,z) Ug(r,z) Cs(r,z)

(6)

The slip coefficient is rather used in the model and correlated directly from gas and catalyst radioactive tracings on a commercial riser. The slip SL(r,z) is represented by the following correlation:

SL(r,z) ) (SLav(z) - SLmin)

(

)

3nsl + 1 × nsl + 1

( () ) 1-

r R

nsl+1/nsl

+ SLmin (7)

Three parameters define the local slip coefficient: nsl, SLmin, and SLav. nsl and SLmin are invariant in the reactor, but SLav is a function of z. Table 2 shows that nsl does not depend strongly on operating conditions for Ug between 3.9 and 6.7 m/s and Fs between 85 and 303 kg‚m-2‚s-1. Thus, nsl is chosen to be the same in the commercial riser and in the cold setup, i.e., 0.35.

Figure 9. Catalyst density in the cold setup at half riser height and close to the elbow.


Figure 10. Slip profiles obtained in the cold setup (Ug ) 5.2 m/s, Fs ) 215 kg‚m-2‚s-1, 2 m upstream of the elbow). Table 2. Slip Coefficients Determined for Various Operating Conditions in the Cold Setup Ug Fs (m/s) (kg‚m-2‚s-1) 3.90 3.90 3.90 3.90 3.90 3.90 5.20 5.20 5.20 5.20 6.70 6.70 6.70 6.70 6.70 6.70

85 85 85 91 111 111 85 135 215 215 170 170 170 303 303 303

position

Nsl

SLav

SLmin

half riser height half riser height 2 m upstream of elbow half riser height half riser height half riser height 2 m upstream of elbow 2 m upstream of elbow half riser height 2 m upstream of elbow half riser height half riser height 2 m upstream of elbow half riser height half riser height 2 m upstream of elbow

0.20 0.19 0.47 0.24 0.62 0.52 0.31 0.50 0.27 0.53 0.13 0.32 0.34 0.26 0.25 0.38

0.36 0.65 0.42 0.50 0.35 0.35 0.64 0.48 0.53 0.46 1.20 0.55 0.58 0.81 0.75 0.63

-1.7 -0.07 -0.29 -0.69 -0.17 -0.24 -0.38 -0.30 -0.43 -0.39 -0.46 -0.14 -0.20 -0.33 -0.51 -0.26

SLmin is always negative in the cold setup, but it was found to be slightly positive or close to zero in commercial risers because of the high density and velocity of the gas. The value of 0.05 was chosen for the model. Radioactive tracings of the gas and the solid in the industrial unit supply information on SLav. It is easy to estimate gas and solid velocities from the first order momentum of the outputs or with an axial dispersed plug-flow model. However, these approaches, often used for processing radioactive tracing data, are inaccurate, because the detector does not measure an average concentration and flows are not plug, as shown above. For instance, Figure 11 shows the results of tracer experiments in a commercial riser. These data are analyzed with a plug-flow model with axial dispersion.

Figure 11. Gas radioactive tracing in a commercial riser; processing with an axial dispersion model.

Detectors are placed at the bottom and the top of the risers. The Peclet number is always found to be large, in this case superior to 30, which suggests a flow close to perfect plug. However, the model never fits well, as


shown in Figure 11, since the front section of the pulses are always steeper than the computed ones. The justification of the tracer profiles found is quite clear once the flow characteristics described above are accounted for: the steep front of the peak corresponds to the apparition of gas with the highest velocity, i.e., gas flowing in the center line of the riser. Then the long tail corresponds rather to gas flowing closer and closer to the wall than to a significant axial dispersion. This interpretation is difficult to quantify because tracer molecules emitting γ-rays are not seen equally by the detector for several reasons: (i) The fastest evolving tracer fractions are underestimated since they remain a lesser time in the vision field of the detector, which receives less counts. (ii) The tracer fractions evolving far from the detector are underestimated because γ-rays are emitted in every direction so that the intensity recorded by the detector is inversely proportional to the squared distance. Moreover, these tracer fractions are even more attenuated by the various material in the riser, particularly by catalyst with its core-annulus structure. The same observations can be done for tracing catalyst, with a small difference: the axially dispersed plugflow model always fits better probably because solid is more backmixed than gas. To reconstruct the flow from a detector output would need the knowledge of the whole solution: velocity profiles and material density mapping, plus absorption characteristics of the riser wall. The detection geometry should also be introduced. To accomplish all this on a routine basis would be unrealistic. Given this, a simpler approach is proposed. Since it is shown that gas and catalyst have the highest velocity at the center line of the riser, this velocity is measured from their arrival time recorded by the detector. By using the Ostwald-de Waele representation, it is then possible to calculate the velocity profile in the radial direction. Moreover, the gas-solid slip coefficient SLax is defined as their velocity ratio at the riser center line derived from the phase’s arrival time, and this is correlated by a polynomial expression to the riser relative height Z.

SLax ) -2.456Z2 + 3.058Z + 0.054

(8)

Notice that the slip coefficient at maximum velocity is small at Z ) 0, is close to 1 at Z ) 0.6, and decreases again to 0.57 at Z ) 1. Combining eqs 7 and 8 yields

SL(r,z) ) (SLax - SLmin)(1 - rnsl+1/nsl) + SLmin

(9)

Thus, the modeling of the gas velocity profile and of the slip coefficient allows one to map the catalyst holdup from the catalyst flux correlation. Description of the Kinetics of Catalytic Cracking Describing the kinetics of catalytic cracking is no mean task: the feedstock is a complex mixture of hydrocarbons and the chemical reactions are numerous. An exact description of these reactions would be very difficult, and the resulting system would be too complex to combine it with hydrodynamics. In a former work, our group used with success a four-lump model (feedstock, gasoline, gas, coke) (Martin et al., 1992b; Pitault et al., 1995). Unfortunately, the kinetic constants

Table 3. Lumps Considered in the Feedstock and the Cracking Products feedstock: C25-C42 ; Teb > 350 °C paraffins naphthenes with 1, 2, or 3 cycles naphthenes with more than 3 cycles aromatic compounds LCO: C13-C24 ; Teb ) 215-350 °C paraffins olefins naphthenes olefinic naphthenes aromatic compounds gasoline: C5-C12 ; Teb) 40-215 °C paraffins olefins naphthenic hydrocarbons olefinic naphthenes aromatic compounds LPG: C3-C4 paraffins olefins fuel gas: H2 C1-C2 coke, Conradson carbon Table 4. Main Types of Reactions Considered β-scission reactions of paraffins: CnH2n+2 f CmH2m + CpH2p+2 of olefins: CnH2n f CmH2m + CpH2p of alkyl chains of aromatic compounds: Ar-CnH2n+1 f Ar-H + CnH2n of naphthenes: CnH2n(naphthene) f CnH2n(olefin) condensation reactions aromatic compound + olefin f aromatic compound (including coke) cyclization reactions CnH2n(olefin) f CnH2n(naphthene) hydrogen-transfer reactions naphthene + olefin f aromatic compound + paraffin other reactions feed components f light gas (H2, C1, C2) Conradson carbon f coke

obtained in a pulsed microactivity test strongly depended on the feedstock used. Here we use a more complex, but also more general, kinetic model developed by Pitault et al. (1994). It comprises 19 chemical lumps and 25 chemical reactions. Tables 3 and 4 present the lumps and the main reactions taken into account. All the catalytic reactions are inhibited by the presence of coke on the catalyst. The results of Pitault et al. (1994) show that one single deactivation factor φ by coking can be used for all reactions, except the reactions leading to formation of coke. For the latter reactions, the deactivation φ decreases faster with coke content. Each reaction rate rAB,P of A + B f P takes the general form rAB,P ) kAB,Ps φPCACB. The details of the kinetic laws proposed may be found in Pitault et al. (1994). Prediction of the Conversion of the Feedstock and of the Yield of the Products The reactions are assumed to begin at the bottom of the riser, where the gas velocity profile is calculated according to the results of the hydrodynamic investigation. Then, for each component of the gas phase and for the coke content of the solid, the following balance equations can be written (assuming the velocities to be vertical):

( )

∂(UgCi) Dr ∂ ∂Ci ) r + Ri ∂z r ∂r ∂r

(10)


Figure 12. Catalyst concentration in a commercial riser: experimental data and model results at two different gas radial dispersion.

and

( )

Fs ∂Cc D′r ∂ ∂Cc ) r + Rc Fp ∂z r ∂r ∂r

(11)

where the production rates of component i and of coke (c) are obtained by algebraic summation of all reactions producing or consuming component i (c):

rij + ∑ ∑ j*i i*j

Ri ) -

cons.

βijrij

(12)

prod.

Herein, βij is a stoichiometric coefficient. The solid holdup s included in each term of eq 12 and needed to estimate the reaction rates is obtained via the gas velocity and the correlation of the slip factor. The gas velocity Ug is obtained by adding all the balance equations of the gas components and by writing that total pressure and temperature remain constant:

∂Ug

RT )

∂z

∑ Ri

Ptot i*c

(13)

This set of equations is an extension of the equations of the model of Martin et al. (1992b) to the more complex kinetic scheme considered here. A combination of this set of parabolic coupled partial differential equations with appropriate boundary conditions (zero concentration gradient on the axis and at the wall) yields the concentrations of the different products at the column outlet. The solution of this system was obtained with the help of a commercial software. Besides the simplifications of the physical and chemical reality which are described in this paper, one must recall some other assumptions underlying these equations: The reactor is supposed to be isothermal, as it is known that endothermic effects of 40-60 °C exist in the plant. Strong mass- and heat-transfer limitations exist at the riser bottom, and they can be more or less attenuated by proper catalyst/feedstock mixing devices.

Figure 13. Effect of the core-annulus structure on conversion at 20 m in a commercial plant and comparison with the model.

The gas convection is only in the vertical direction. It is fixed at the riser’s bottom, and it varies then only because of molar expansion due to cracking. This yields, at z > 0, nonzero gas velocity at the wall, which is certainly not realistic. The solid flux is considered to be constant along the vertical direction. The catalyst local holdup develops according to the evolution of the gas velocity and the imposed slip relation, so that the core-annulus structure vanishes somewhat when along the riser. This is because the gas velocity increases more at the wall since there is locally a smaller WHSV. Fortunately, this evolution corresponds to the physical reality: it is known that the core-annulus structure is less pronounced when the superficial gas velocity increases. Contrary to the observations in Figure 9, this model does not predict that core-annulus structure tends to be formed at the riser exit so that its effect on conversion is underestimated. Results of the Model and On-Site Validation In summary, the model is built from hydrodynamic correlations obtained on the cold setup, except for the definition of the slip coefficient SLax in the center line, which is fitted to the riser height by radioactivity in the commercial plant. The kinetic model is set up from laboratory experiments with a microreactor. This model was first checked by comparison with plant data where maps of catalyst density by γ-ray


Figure 14. Axial yield profiles: plant data and model.

tomography were available. Here, the C/O, i.e., the weight ratio of catalyst flux to oil flux, is 5.5. Figure 12 represents the radial profile of catalyst concentration some meters above the feedstock injection. The model was run with radial dispersions of 10-3 and 3 × 10-2 m2/s. This variation does not change significantly the catalyst concentration in the dilute area, but the dense zone close to the wall is affected. There, a larger dispersion provokes more radial exchange of chemical species, so that unconverted feedstock replaces less reactive cracked products whose production is enhanced by the small local space velocity. This allows a larger increase of gas velocity at the wall and thus a smaller catalyst density. The model predicts well the catalyst concentration profile, although it is somewhat underestimated. The model predicts also that radial oil product profiles are not even and that the particular hydrodynamics induce a loss of conversion and selectivity (Martin et al., 1992b). This fact was further checked in a commercial plant at 2/3 riser height, where hydrodynamics are well established and effects of feedstock jets are supposed to vanish. Several samples of cracked oil were taken on the riser diameter and analyzed. The radial profile of conversion to gas and gasoline is shown in Figure 13. Obviously, the effect of the core-annulus structure of the catalyst flow combined with the Ostwald-de Waele type flow of the gas is responsible of this profile. The model was used to fit the data. However, the conversion level was hardly attained when using the real C/O of the plant. This lack of conversion was attributed to the underestimation of catalyst density shown in Figure 12 and also to the fact that the isothermal kinetic model is set at 530 °C, a temperature below the average temperature of this endothermic conversion. Therefore, the model was fitted on Figure 13 data by optimization of the radial gas dispersion coefficient and on the C/O. The model results of Figure 13 are obtained with a

dispersion of 0.035 m2/s, a value very close to the one obtained from helium tracing (Figure 7). The C/O is adjusted to 11.5. Using the same data, the model results were also compared to an other commercial runs where several samples were taken along the riser center line. The agreement with conversion and gasoline yield of the model at the center line is very good, as shown in Figure 14. The conversion and gasoline yield averaged on the riser cross section are also computed and shown. The radial yield gradients are large between 4 and 10 m and decrease somewhat at the riser outlet because of chemical kinetic limitations and of radial mixing. However, the differences still remain significant. A purpose of such a model is to evaluate how fluid dynamics affect yields and on this basis to be able to improve reactor performance, via development of internals or downflow reactors. Therefore, the influence of gas radial dispersion was studied: at high dispersion the reactor behaves like a perfect plug flow, whereas at low dispersion hydrodynamic effects are enhanced. The results are shown in Figure 15. Conversion and gasoline yield can be improved by ca. 3 wt % if the reactor behaves as an ideal plug flow. This is not negligible and may justify further developments. Products quality becomes more and more important because of environmental considerations. The model teaches us whether riser hydrodynamics change products quality. This is shown in Figure 16, where the gasoline analysis is plotted versus the gas radial dispersion. The four chemical lumps of the gasoline are paraffins, olefins, naphthenes (alicyclic hydrocarbons), and aromatics. When conversion increases, slow hydrogen transfer reactions take place to yield paraffins (mainly branched) and aromatics from olefins and naphthenes. These reactions occur between every lump of the oil mixture (feedstock, light gas oil, gasoline, gases) while cracking is progressing. Small changes are


Figure 15. Effect of gas radial dispersion on conversion and gasoline yield.

Figure 16. Influence of gas radial dispersion on gasoline analysis.

visible in the operating window of gas radial dispersion. A reactor flow pattern closer to the plug flow would enhance the hydrogen-transfer reactions, increasing the yields of aromatics and isoparaffins. The gasoline analysis would be slightly affected, leading to a small increase of motor octane number. This quality change is believed to be not important enough to justify reactor design improvements. There are other effective ways to optimize products quality as in the case of changes of catalyst formulation. Conclusions FCC riser reactors can be modeled in two dimensions by a vertical flow with no backmixing for gases and catalyst and radial profiles of velocity and concentration. A single radial dispersion coefficient is able to account

for the gas behavior in the whole cross-sectional area. This coefficient has been measured directly on a commercial plant. The mixing properties of the catalyst are not important in this process because it reaches very quickly its final coke content and activity. A detailed chemical kinetic model developed from laboratory data provides accurate information on products yields and quality. The comparison of the model to analyses from sampling in commercial plants confirms its validity. This allows one to estimate the loss of conversion and gasoline yield to 3 wt % when the riser is compared to a plug-flow reactor, and this shows that products quality is not significantly affected. Although this model of FCC riser is probably one of the most complete ever published, it still ignores many effects, including (i) feedstock biphasic jets at the riser


bottom, (ii) reinforcement of the core-annulus structure upstream of the elbows (thus, the yield effects are probably slightly underestimated), and (iii) thermal effects. Acknowledgment The authors thank the company ELF ANTAR FRANCE for permission to publish. Collaborators like P. Turlier, R. Barberet, and D. Gadolet are friendly and gratefully acknowledged for their contribution. Nomenclature C ) tracer gas concentration, mol/m3 (eq 2) Cs ) solid concentration, kg/m3 Dr, Dr′ ) radial dispersion coefficients of the gas and of the solid, m2/s Fs ) solid flux, kg‚m-2‚s-1 n ) consistency index in an Ostwald-de Waele type law (eq 1) nsl ) index in eq 7 Qtot ) volume flow rate of the gas, m3/s r ) radial coordinate, m rij ) reaction rate from component i to component j, mol‚m-3‚s-1 R ) radius of the riser, m R ) ideal gas constant, J‚mol-1‚K-1 (eq 13) Ri ) production rate of component i, mol‚m-3‚s-1 SL ) slip coefficient (eq 7) SLax ) slip coefficient on the column axis Ug ) gas velocity, mean gas velocity, m/s z ) axial coordinate, m Z ) dimensionless axial coordinate Greek Letters βij ) stoichiometric coefficient ) porosity s ) solid holdup Fp ) particle density, kg/m3 Indices av ) average exp ) experimental value max ) maximum min ) minimum

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Received for review June 17, 1997 Revised manuscript received September 9, 1997 Accepted September 11, 1997X IE970432R

X Abstract published in Advance ACS Abstracts, October 15, 1997.

Hydrodynamics of Riser Units and Their Impact on FCC Operation

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