Gas Mixing in the Reactor Section of Fluid Cokers - American

Ind. Eng. Chem. Res. 2005, 44, 6067-6074

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Gas Mixing in the Reactor Section of Fluid Cokers Xuqi Song,† John R. Grace,*,† Hsiaotao Bi,† C. Jim Lim,† Edward Chan,‡ Brian Knapper,‡ and Craig A. McKnight‡ Fluidization Research Centre, Department of Chemical and Biological Engineering, University of British Columbia, 2216 Main Mall, Vancouver V6T 1Z4, Canada, and Syncrude Research Centre, 9421 17th Avenue, Edmonton, Alberta T6N 1H4, Canada

The gas flow structure in the reactor section of a fluid coker cold model was investigated using helium as the tracer gas. Gas residence time distributions indicate intensive gas mixing. The mean residence time from the feed nozzles decreased when the same flow of air was distributed from three rings of nozzles rather than six. When steady state measurements were carried out by injecting helium through a feed ring in the middle of the reactor section, the downstream radial profile of helium concentration formed a bell-shaped curve above the jet region near the tracer injection level and became nearly uniform with increasing height due to radial dispersion. Upstream, the helium concentration was higher in the wall region because of descending solids. The core-annulus flow structure disappeared toward the bottom of the reactor section. Introduction Fluid cokers are noncatalytic fluidized bed reactors used to convert heavy hydrocarbons into lighter liquid products, gases, and solid coke. The thermal cracking process is carried out in large fluidized beds, where hot coke particles carry the heat required for the endothermic reactions and collect solid byproducts on their surfaces. The hydrocarbon feed is injected through horizontal nozzles arranged in rows, and the vapor products tend to channel up the center of the bed, surrounded by dense, descending particles. Hydrodynamics in the reactor section, such as the voidage distribution and solids flow structure, have been studied in two geometrically and dynamically scaled cold models.1-3 Gas mixing behavior is also important with respect to the overall performance of cokers in achieving high conversion and selectivity. Backmixing is undesirable as it leads to increased production of dry gas and coke byproducts. From previous studies,4-6 it is known that gas backmixing in fluidized beds is promoted by the downflow of solids near the wall. The solids flow structure in fluid cokers resembles circulating fluidized beds in that a relatively dense annular region, in which the time-averaged flow is downward, surrounds a more dilute, upward-flow core region.2,3 Gas elements are dragged downward in the dense annular region if the absolute downward solids velocity exceeds the relative velocity between the gas and the particles. The internal circulation of solids in fluid cokers is influenced by the diverging cross-sectional area as the gas rises and by the array of horizontal feed nozzles spraying vaporizing hydrocarbon feed toward the center of the bed.2,7 Most gas mixing studies have been carried out by injecting tracer gas at a point in the fluidized bed, with the rest of the gas introduced through a distributor at the bottom of the bed. There is very little work on the effect of geometries such as that encountered in fluid * To whom correspondence should be addressed. Tel: (604) 822-3121. Fax: (604) 822-6003. E-mail: [email protected]. † University of British Columbia. ‡ Syncrude Research Centre.

cokers where the walls are not vertical and where gas is introduced at multiple injection points. In this study, gas residence time distributions and steady state gas mixing behaviors were investigated in a cold model fluid coker. Experimental Equipment and Methodology The experiments were carried out in a fully threedimensional cold model fluidization column, geometrically scaled down by a factor of ∼20 from two commercial fluid cokers operated by Syncrude Canada Limited in Fort McMurray, Alberta, Canada. The setup is shown in Figure 1. The facility includes a reactor section and a stripper section on the coker side, an external riser to return solids removed from the stripper back to the top of the reactor, and a collection system to capture entrained particles. The reactor and stripper sections were designed to be geometrically similar to the corresponding sections of two commercial fluid cokers.3 The net downward solids flow rate was typically 3.4 kg/s, controlled by a pinch valve in the standpipe. The solid particles in this study were FCC particles, with a mean diameter of 99 µm and an apparent density of 1700 kg/m3. Air was the fluidizing gas, and the bed was operated at a somewhat elevated pressure, 200 kPa, to maintain the same ratio of solids density to gas density as in the commercial units. The reactor section was axisymmetric and tapered, expanding from a diameter of 308 mm at the bottom to 485 mm at the top. Despite the tapered geometry, the superficial gas velocity increased with height up the reactor section, since gas was fed in increments through six injection rings. Each ring contained 10-18 nozzles to achieve relatively uniform injection around the outer perimeter. There were six feed rings, equally spaced at intervals of 108 mm. Three levels of nozzles in the lower region supplied the stripper with stripping gas and provided the fluidization air needed in this section. The nozzles in the top row were sparging nozzles, whereas those in the bottom two rows of the cone section were stripping nozzles. An additional ring of nozzles was installed between the reactor and the stripping section to simulate the high velocity steam attrition jets

10.1021/ie040230c CCC: $30.25 © 2005 American Chemical Society Published on Web 02/17/2005

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Ind. Eng. Chem. Res., Vol. 44, No. 16, 2005

Figure 1. Schematic of pressurized fully cylindrical cold model fluid coker experimental facility.

used to maintain a stable particle size distribution in the commercial fluid cokers. The heights of the feed rings and other gas inlets are shown in Figure 2. Above the surface of the fluidized bed, the cold model column expands beyond the geometrically similar dimensions in order to assist in disengaging particles in the freeboard zone. Equalities of the Archimedes number, Reynolds number, solids-to-gas density ratio, and solids-to-gas mass flow ratio were established to provide dynamic similarities in the reactor section of the cold model with the two commercial units. The reactor sections in the cold model and commercial units operated in the bubbling fluidization flow regime with Uf/Uc ) 0.9, with Uc predicted by a correlation based on differential pressure fluctuations.8 Additional details of the experimental facility are provided elsewhere.3 Gas mixing in the reactor section was studied by introducing helium tracer gas into one of the rings of feed nozzles, while the tracer concentration was measured at several levels upstream and downstream of the injection level, as well as at the center of the top exit. Helium was chosen as the tracer because it is inert and non-adsorbing on the particles. FCC was employed as

Figure 2. Axial profile of superficial gas velocity in reactor section of cold model.

the catalyst to facilitate dynamic similarity with the coke particles of the commercial reactors. To enable

Ind. Eng. Chem. Res., Vol. 44, No. 16, 2005 6069

Figure 3. Helium tracer gas injection system.

proper mixing of the helium tracer with conveying air upstream of the nozzles and to achieve a uniform distribution over the cross-section, the tracer was added into the manifold air header, as shown in Figure 3. Polyethylene tubes of 12.7 mm OD of the same length (3 m) for each ring connected each nozzle to the manifold. Nozzles were normally inserted 29 mm into the reactor. The concentration of helium at each of the pre-chosen positions in the reactor model was detected by a thermal conductivity cell (TCD). The TCD was calibrated for different known concentrations of helium tracer gas (0-1% helium by volume). Calibration was carried out for the same operating pressure as in the experiments and covered the entire range of helium concentrations encountered during operation. The calibration procedure was repeated for each run performed on different days and found to be reproducible, with less than 5% error. The output signals from the TCDs were converted to helium concentrations using the pre-calibrated linear equation. The tracer injection rate was 6.5 × 10-4 Nm3/s. The sampling/detector system consisted of a 15 µm sintered filter welded onto a 4.6 mm id stainless steel sampling tube. A needle valve was installed before the TCD in the sampling line to suppress pressure fluctuations. Unsteady State Measurements. To obtain gas residence time distribution (RTD) curves in the reactor section, dispersion in the injection/sampling system had to be minimized and characterized. The detector/sampling system RTD was measured by sampling immediately downstream of the feed nozzles, and the mean residence time was determined experimentally to be 2.38 s. As shown in Figure 4, the tracer injector/ detector system gave a sharp step response in the helium concentration. An electronic signal initiated both helium flow into the manifold and recording of the amplified TCD output by a computer, using a customdesigned program at a sampling speed of 10 Hz. Positive step changes of helium injection were used in the experimental determinations of dispersion. Steady State Measurements. In some runs, helium tracer was injected continuously and steadily through all 16 nozzles of feed ring 3 (see Figure 1) to investigate gas mixing in the reactor section. The tracer injection

Figure 4. Detector and detector/coker response to step increase in tracer injection from different positions. Uf ) 0.74 m/s, Us ) 0.25 m/s, and Gsf ) 18.6 kg/m2s.

rate was 2.5 vol % of the total airflow from feed ring 3. The detector probe was positioned at four levels downstream and five levels upstream of the injection level and could be traversed from the wall to the axis while the system was in operation. Data were recorded at a frequency of 10 Hz for sampling periods of 70 s. Results and Discussion The superficial gas velocity profile with six active feed rings, shown in Figure 2, was similar to that in the commercial units, with the superficial gas velocity decreasing with height between successive feed nozzle rings because of the diverging cross-section of the reactor. The reactor section operated in the bubbling fluidization flow regime, whereas the superficial gas velocity at the top of the bed (Uf ) 0.74 m/s) was close to the transition velocity from the bubbling to the turbulent fluidization flow regime (Uc ) 0.80 m/s). Uc is predicted by a correlation based on differential pressure fluctuations:8

Rec ) 1.243Ar0.447

(1)

Experiments were carried out for two operating conditions: gas feed from all six rings and gas feed only from the top three rings while keeping other param-

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Table 1. Residence Time of Tracer in Reactor Section

feeding condition feed from all six rings feed from top three rings

tracer injection level

mean residence time from tracer injection ring to exit, ht (s)

dimensionless standard deviation, σ/th

feed ring 1 feed ring 3 feed ring 6 feed ring 1 feed ring 3

5.35 5.37 5.23 5.01 4.80

0.30 0.32 0.33 0.28 0.31

eters, such as solids circulation rate and total gas feed rate, unchanged. The corresponding axial distributions of cross-sectional average voidage in the reactor section measured by an optical fiber probe (see insert in top left corner of Figure 2) change slightly with height, no doubt influenced by the axial distribution of superficial gas velocity. Unsteady State Measurements. Gas RTD experiments were carried out using positive step tracer inputs into one ring of the feed nozzles. Given the linear relationship between the TCD output and the helium concentration, each response curve can be easily converted to the cumulative distribution function F(t) by

F(t) )

V(t) - V0 V∞ - V0

(2)

Typical experimental F curves are shown in Figure 4. Because the experiments were conducted with the tracer concentration detected at the column outlet, gas RTDs determined from these data include dispersion in the freeboard region. Experimental data were next smoothed to produce F(t) curves. Residence time distributions were then derived by differentiation, i.e.

E(t) )

dF(t) dt

(3)

The mean residence time, ht, and the variance, σ2, of the distribution were then calculated from:

∫0∞t ‚ E(t) dt

(4)

∫0∞(t - ht)2 ‚ E(t) dt

(5)

ht ) σ2 )

Figure 5. Response curves for feed from six and three rings of nozzles. Tracer injected from feed ring 3. Uf ) 0.74 m/s, Us ) 0.25 m/s, and Gsf ) 18.6 kg/m2s.

The system transfer characteristics were then estimated from

ht (reactor) ) ht (reactor + detector) - ht (detector) (6) σ2(reactor) ) σ2(reactor + detector) - σ2(detector) (7) The results are shown in Table 1. Note that the calculated reactor mean residence times are crude estimates because the volume of the freeboard is about six times the void space in the dense fluidized bed in the reactor section above the tracer injection level. This ratio becomes even greater when injecting tracer gas from the upper feed nozzles. The difference in the measured total mean residence time when injecting tracer gas into different feed rings was small, indicating intensive gas mixing in the reactor. Experimental F curves for different feed conditions with tracer injected through feed ring 3 are compared

Figure 6. Helium radial steady state concentration profile in reactor section 54 mm above feed ring 6 where tracer was introduced. Ug ) 0.29 m/s and Gs ) 31.6 kg/m2s.

in Figure 5. As shown in Table 1, the mean residence time from feed rings 1 and 3 was shorter when the same volumetric flow of air was injected only from the top three rings, since the jets penetrate deeper into the bed, and air from the feed nozzles burps up the center of the column as large bubbles. This might be helpful in optimizing and improving the performance of the fluid coker, as shorter residence times of the feed may result in higher liquid yields for residual oil cracking. Of course, other factors such as the solid interaction with the feed jets also affect the reaction. Steady State Measurements. Steady state tracer injection was applied to evaluate radial mixing and backmixing from tracer concentrations monitored at different radial positions downstream and upstream of the injection level. A typical radial concentration profile of the tracer gas 54 mm above the helium injection level is shown in Figure 6. The jet penetration depth, L, defined as the horizontal distance from the tip of the nozzle to the time mean position of the end of the jet region, was estimated from the semiempirical correlation of Merry:9

[

F0u02 L + 4.5 ) 5.25 d0 (1 - )Fpgdp

]()() 0.4

Fg Fp

0.2

dp d0

0.2

(8)

This equation has previously been found1,3,10,11 to give good predictions for our system. The shape and penetration of the jets are also shown in Figure 6. The radial profile of helium concentration is approximately


Figure 7. Helium radial steady state concentration profiles in reactor section 54 mm above feed ring 3 where tracer was introduced for feed from six and three rings of nozzles.

axisymmetric as the tracer gas is injected from 10 uniformly distributed nozzles in the feed ring. However, the measured helium concentrations in the left tail (r/R ) -1 to -0.6) are appreciably higher than in the right tail (r/R ) 0.6 to 1). The probe is inserted from the right side (r/R ) 1). Hence, the results on the right side of the reactor are less intrusive and likely to be more accurate than on the left. In what follows, we plot only the data from the right side. The relatively high concentration of tracer above the jets can be attributed to the sampling probe mostly capturing bubbles, generated from the helium-containing jets. The helium concentration decreases toward the axis of the column indicating that the feed jets did not reach the axis. It also can be seen from the radial profile of helium concentration that Merry’s correlation9 provides reasonable agreement for the experimental conditions of this study. Radial concentration profiles 54 mm above feed ring 3 are compared in Figure 7 for the same total feed rate introduced through six, or only the top three, feed rings. Jets penetrate further when the same total air flow rate is redistributed only into the top three rings, causing the concentration of helium to be higher at the center of the column. Overall distributions of helium concentration in the reactor section, both upstream and downstream of the helium injection level (feed ring 3), are shown in Figures 8 and 9 for six and three active feed rings, respectively. Nozzle tip and jet end positions, calculated from eq 8, are also shown. The corresponding axial distributions of superficial gas velocity appear in Figure 3. The helium concentrations above the jet region are higher than at other locations within a limited distance downstream of the injection level. The nonuniformity of concentration across the column decreases with distance above the tracer injection ring, due at least in part to radial gas mixing. In the fully developed mixing zone well above the tracer injection level, the radial concentration profiles become nearly uniform. In Figure 9, the

Figure 8. Radial and axial distribution of helium concentration in reactor section. Feed from six rings. Uf ) 0.74 m/s, Us ) 0.25 m/s, and Gsf ) 18.6 kg/m2s. Tracer injected with air via feed ring 3. Jet penetration estimated from eq 8. Horizontal bars show average concentrations in core and annular regions.

radial profile of the helium concentration does not become uniform, even at the upper level of the bed, because big bubbles generated from the jet rise quickly at the center of the column. This also indicates less radial gas mixing and lower gas-solids interaction efficiency for this feed configuration. The upstream tracer concentrations are higher near the wall than in the core of the column. The radial concentration profiles indicate considerable axial dispersion of gas near the wall due to downflow of solids there. This solids downward flow was measured by an optical fiber velocity probe in the same equipment, as reported by Song et al.3 Downflow of gas can occur when the downward velocity of solids exceeds the relative interstitial gas velocity in the dense phase.12 Tracer gas, together with the solid particles, is entrained by the feed jets from the downward flowing annular region into the upward flowing core region. Note that the tracer gas penetrates downward to the top of the stripper section due to the vigorous internal circulation of gas and solids in the reactor section. The radial gas velocity profiles are influenced by the radial profiles of voidage and gas velocity. A simple relationship between dense bed cross-sectional average

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Figure 10. Axial distribution of cross-sectional average voidage in reactor section from optical probe measurements and from eq 9. Feed from all six rings. Uf ) 0.74 m/s, Us ) 0.25 m/s, and Gsf ) 18.6 kg/m2s.

Figure 9. Overall distribution of helium concentration in reactor section. Feed from three rings. Uf ) 0.74 m/s, Us ) 0.25 m/s, and Gsf ) 18.6 kg/m2s. Tracer injected with air via feed ring 3. Jet penetration estimated from eq 8.

voidage and superficial gas velocity proposed by King13 is

j )

Ug + 1 Ug + 2

(9)

Radial voidage profiles, (r), were measured in our equipment by an optical fiber probe, as described in detail elsewhere.3 The axial distribution of measured cross-sectional average voidage is compared with that predicted by eq 9 in Figure 10. Equation 9 is seen to give reasonable agreement with the experimental results. On the basis of the hypothesis that the relationship between the local superficial gas velocity, U(r), and the local voidage, (r), follows the same functional relationship as eq 9, U(r) can then be estimated from the measured local voidage data by

U(r) )

2(r) - 1 1 - (r)

Figure 11. Radial profiles of local superficial gas velocities in reactor section calculated from eq 10. Feed from all six rings. Uf ) 0.74 m/s, Us ) 0.25 m/s, and Gsf ) 18.6 kg/m2s.

it was confirmed that particles tend to rise in the central core, while descending in the outer annular region. The radii of the core region at different heights were based on the positions where the time mean particle velocity passed through zero. The average dimensionless radial position (rc/R) of this boundary is 0.77 in the reactor section. The axial profiles of the average helium concentration in the core and annular regions are shown in Figure 12 for a case with six feed rings operating. The average helium concentrations in the annulus and h c, are based on core regions, C h a and C

∫rRC(r)U(r)r dr ∫0r C(r)U(r)r dr Ca ) ; Cc ) ∫rRU(r)r dr ∫0r U(r)r dr c

c

(10)

Figure 11 shows the radial profiles of local superficial gas velocity for five heights. Changes of U(r) with height are predicted to be more significant in the core region than near the wall. In our previous study,3 by measuring the particle velocity in the reactor section with an optical fiber probe,

c

(11)

c

Average helium concentrations in the annular region are higher than in the core region upstream, except near the bottom of the reactor section, where the radial profile of helium concentration becomes nearly uniform. Measurements of particle velocity3 also indicate disappearance of the core-annulus structure at this level.


Figure 12. Axial profiles of average helium concentrations in core and annular regions in reactor section upstream of injection level (feed ring 3).

Gas backmixing in the annulus region differs fundamentally from that in the core region. Interchange of gas and solids between the core and the annulus regions is promoted by the horizontal jets. Given these factors and the variation in superficial gas velocity in the reactor section due to the intermediate feed rings and the tapered column, a homogeneous one-dimensional model is unable to describe the dispersion in the reactor section. A two-dimensional model is needed to describe the gas mixing behavior in the reactor section of fluid cokers. Such a model is currently under development. Conclusions The flow behavior of gas in the reactor section of a cold model pilot scale column, geometrically and dynamically scaled to represent two commercial fluid cokers, was studied using unsteady and steady state tracer injection. The mean residence time of gas from the same feed ring level became shorter when feed was injected from the top three rings rather than from all six rings. Radial profiles of helium concentration showed maxima above the position of maximum jet penetration and then became uniform with increasing height due to radial dispersion. It required a longer distance above the tracer injection level for the helium concentration to become uniform when feed entered only from the top three rings. Upstream, the helium concentration was highest near the wall because of mainly downward solids flow there. Gas was dragged downward by the solids descending in the wall region and then rose rapidly in the core region. Average helium concentrations in the annular and core regions differed significantly, with the boundary defined as the radial position at which the time mean particle velocity passed through zero. The core-annulus flow structure disappeared at the bottom of the reactor section. Acknowledgment We are grateful to Syncrude Canada Limited for sponsoring this work and for permission to publish the results. We are also grateful to Chevron Canada Ltd. for providing the FCC particles. Nomenclature Ar ) Archimedes number, Ar ) Fg(Fp - Fg)dp3 g/µ C ) concentration of helium, Ca ) concentration of helium in annular region, -

Cc ) concentration of helium in core region, d0 ) inner diameter of nozzle, m dp ) particle diameter, m E(t) ) exit age distribution function, s-1 F(t) ) cumulative distribution function, Gs ) local net solids circulation flux corrected for crosssectional area, kg/m2s Gsf ) net solids circulation flux based on cross-sectional area at top of reactor, kg/m2s L ) horizontal jet penetration, m R ) radius of column, m Rec ) Reynolds number, Rec ) FgUcdp/µ r ) radial distance from center, m rc ) core radius, m t ) time, s ht ) mean residence time, s u0 ) feed nozzle jet velocity, m/s Uc ) transition velocity at which standard deviation of pressure fluctuations reaches a maximum, m/s Uf ) superficial gas velocity at dense phase bed upper surface, m/s Ug ) superficial gas velocity at a given level, m/s Us ) superficial gas velocity at top of stripper section, m/s U(r) ) local superficial gas velocity, m/s V0 ) average TCD output before step change in helium concentration, volts V∞ ) average TCD output long after step change in helium concentration, volts V(t) ) TCD output at time t after step change in helium injection, volts x ) vertical distance from tracer gas injection level, m Z ) height coordinate measured from top of highest stripper baffle, m Z* ) dimensionless height coordinate (Z divided by total height of reactor section), Greek Letters (r) ) local voidage, µ ) gas viscosity, Pa s Fo ) gas density at nozzle tip, kg/m3 Fg ) gas density in fluidized bed, kg/m3 Fp ) particle density, kg/m3 σ ) spread in residence times, s

Literature Cited (1) Knapper, B. A. Experimental studies on the hydrodynamics of fluid bed cokers. M. Sc. Thesis, University of Saskatchewan, Saskatoon, Canada, 2000. (2) Knapper, B.; Berruti, F.; Grace, J. R.; Bi, H. T.; Lim, C. J. Hydrodynamic characterization of fluid bed cokers. In Circulating Fluidized Bed Technology VII; Grace, J. R., Zhu, J., de Lasa, H., Eds.; C. S. Ch. E.: Ottawa: 2002, p 263. (3) Song, X. Q.; Bi, H. T.; Lim, C. J.; Grace, J. R. Hydrodynamics of the reactor section in fluid cokers. Powder Technol. 2004, 147, 126. (4) van Deemter, J. J. Mixing patterns in large-scale fluidized beds. In Fluidization; Grace, J. R., Matsen, J. M., Eds.; Plenum: New York, 1980; p 69. (5) Li, J. H.; Weinstein, H. An experimental comparison of gas backmixing in fluidized beds across the regime spectrum. Chem. Eng. Sci. 1989, 44, 1697. (6) Namkung, W.; Kim, S. D. Gas backmixing in a circulating fluidized bed. Powder Technol. 1998, 99, 70. (7) Matsen, J. M. Scale-up of fluidized bed process: Principle and practice. Powder Technol. 1996, 88, 237. (8) Bi, H. T.; Grace, J. R. Effects of measurement method on velocities used to demarcate the transition to turbulent fluidization. Chem. Eng. J. 1995, 57, 261. (9) Merry, J. M. D. Penetration of a horizontal gas jet into a fluidized bed. Trans. Inst. Chem. Eng. 1971, 49, 189.

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(10) Copan, J. Macroscopic modeling of a fluid bed coker and experimental studies of one- and two-phase feed jets. M. Sc. Thesis, University of Saskatchewan, Saskatoon, Canada, 1999. (11) Donald, A.; Bi, H. T.; Grace, J. R.; Lim, C. J. Penetration of single and multiple horizontal jets into fluidized beds. In Fluidization XI; Arena, U., Chirone, R., Miccio, M., Salatino, P., Eds.; Engineering Conferences International: New York, 2004; p 171. (12) Stephens, G. K.; Sinclair, R. J.; Potter, O. E. Gas exchange between bubbles and dense phase in a fluidized bed. Powder Technol. 1967, 1, 157.

(13) King, D. F. Estimation of dense bed voidage in fast and slow fluidized beds of FCC catalyst. In Fluidization VI; Grace, J. R., Shemilt, L. W., Bergougnou, M. A., Eds.; Engineering Foundation: New York, 1989; p 1.

Received for review August 31, 2004 Revised manuscript received December 8, 2004 Accepted December 10, 2004 IE040230C

Gas Mixing in the Reactor Section of Fluid Cokers - American

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