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Hydrodynamics of the reactor section in fluid cokers
Powder Technology 147 (2004) 126 – 136 www.elsevier.com/locate/powtec
Hydrodynamics of the reactor section in fluid cokers Xuqi Songa, Hsiaotao Bia, C. Jim Lima, John R. Gracea,*, Edward Chanb, Brian Knapperb, Craig McKnightb a
Fluidization Research Centre, Department of Chemical and Biological Engineering, University of British Columbia, 2216 Main Mall, Vancouver, British Columbia, Canada V6T 1Z4 b Syncrude Research Centre, 9421-17th Avenue, Edmonton, Alberta, Canada T6N 1H4 Received 6 January 2004; received in revised form 18 August 2004; accepted 17 September 2004 Available online 6 November 2004
Abstract The hydrodynamics of fluid cokers were studied in a pressurized fully cylindrical cold model scaled by matching key dimensionless groups so that the reactor sections of the commercial and laboratory units were geometrically and dynamically similar. Voidage distributions and solids flow structure were investigated using optical fibre probes in a fluidized bed of FCC. The results show a relatively dense annular region, in which the time-average flow is downwards, surrounding a more dilute, upward flowing core region. However, the radial profiles of voidage indicate that the voidage increases gradually from the wall to the axis of the column, rather than showing a sharp transition between an annular and core region. Particles in the descending outer region are entrained by the feed jets into the core region. The radial position of the boundary between upward and downward net solids flow did not change significantly when the superficial gas velocity was varied. D 2004 Elsevier B.V. All rights reserved. Keywords: Fluidization; Fluid coker; Hydrodynamics; Voidage
1. Introduction Fluid coking is a non-catalytic pyrolysis process used to convert high-molecular-weight hydrocarbons to lower boiling point hydrocarbons, gases and coke. The thermal cracking process is carried out in large fluidized beds, where hot coke particles, introduced from above, carry the heat required for the endothermic reactions and collect solid byproducts on their surfaces. The hydrocarbon feed is injected through several rows of horizontal nozzles and the vapor products tend to migrate up the centre of the bed, surrounded by dense, descending particles [1]. Particles descend through the reactor and then through a stripper, in which a countercurrent flow of steam strips liquid hydrocarbon from the particle surface. The coke particles are then circulated into a fluidized bed burner for re-heating
* Corresponding author. Tel.: +1 604 822 3121; fax: +1 604 822 6003. E-mail address: [email protected] (J.R. Grace). 0032-5910/$ - see front matter D 2004 Elsevier B.V. All rights reserved. doi:10.1016/j.powtec.2004.09.033
prior to re-entering the fluid coker at an increased temperature [2]. After being steam-atomized and injected through feed nozzles, the heavy hydrocarbon feed remains as a thin liquid film on coke particles within the reactor until the feed has fully reacted to form vapour products or coke byproduct. Thin films of feed on coke provide maximum liquid yield [3]. The distribution of liquid feed over the coke solids is strongly dependent on the mode in which the spray from the feed nozzles interacts with the surrounding particles. In order to optimize and improve the performance, it is important to understand the radial distribution of solids concentration and the particle flow structure within the reactor. To improve the understanding of the hydrodynamics in the coker, two cold flow models (a semi-circular Plexiglas column and a fully cylindrical steel column) were designed, constructed and commissioned at the University of British Columbia. The semi-circular cold model has been used primarily to study the hydrodynamics of the stripper section
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to guide the operation and improvement in baffle configuration and solids circulation. Several operational problems, such as fouling and flooding have been investigated in this equipment [4,5]. Measurements of key characteristic parameters in the reactor section were also made for various operating conditions in the semicircular cold model [1,6]. However, the semi-cylindrical column was designed to provide similarity with commercial cokers in the stripper section, not the reactor section. Moreover, being semicylindrical, the column introduces wall and corner effects which make it difficult to apply the results to the fully cylindrical commercial cokers. Indeed the flat face of the semi-cylindrical column was found to have a significant effect on the cross-sectional distribution of voidage and axial solids flow patterns [6]. A pressurized cold model facility was therefore designed, constructed and commissioned to examine the hydrodynamic behaviour in the reactor section of a column whose cross-section was fully circular.
2. Experimental An overall schematic of the fully cylindrical cold-model facility is shown in Fig. 1. All components were fabricated of steel. These include a reactor section and a stripper section on the coker side, a riser for external solids circulation and a collection system for entrained particles. The equipment was designed so that its reactor section would be both geometri-
Fig. 1. Schematic of UBC pressurized fully cylindrical cold model of Syncrude Cokers.
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Fig. 2. Cumulative mass fractions for FCC and Fluid Coke particles.
cally and dynamically similar to those of two commercial fluid cokers operated by Syncrude Canada Limited in Fort McMurray, Alberta, Canada. Scaling, based on dimensionless analysis [7] was achieved by matching the important dimensionless hydrodynamic parameters Ar; Fr;
qp Gs ; ; PSD; H=D qg qg Uf
ð1Þ
Due to the large geometric scale-down factor and the fact that the particles used in the commercial cokers are already quite small, it was not possible to keep the ratio of D/d p constant. However, D/d pNN1 in both cases, so that D/d p matching can be relaxed, so long as all gaps through which the particles flow in the reactor and stripper sections are wider than 50d p~. Above the bed surface the cold model column expands beyond the geometric similarity dimensions in order to assist in disengaging particles in the freeboard zone. Plots of the cumulative mass fraction for the FCC and fluid coke particles are shown in Fig. 2. FCC particles were chosen for this study because optical fibre probes can be used to investigate hydrodynamics of these nearly white particles in the cold model, whereas this is impossible with the black coke. Other properties appear in Table 1. The dimensionless PSDs are quite similar for the FCC and coke particles, as required by the scaling criteria. The results of scaling calculations based on the superficial gas velocity at the surface of the dense bed in the reactor section are summarized in Table 1. The bbase conditionsQ referred to hereafter are the operating conditions (with all 6 feed rings operating, U s=0.25 m/s; U f =0.74 m/s; G sf=18.6 kg/m2 s) having dynamic similitude with the commercial units under normal operation. Experiments by Farrell et al. [8] showed that the solid-togas density ratio must be matched to ensure accurate scaling of the hydrodynamics in bubbling or turbulent fluidized beds. To maintain the same density ratio, q p/q g, as in the commercial units, the cold model was operated at a slightly elevated pressure, 200 kPa, at the bottom of the reactor section. This section is axisymmetric and tapered, expand-
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Table 1 Summary of dimensional similitude based on reactor section
Gas
Temperature, 8C Pressure, kPa Gas density, kg/m3 Gas viscosity, Pa s Type of particles Particle density, kg/m3 Mean particle diameter, Am Geldart powder groupa Superficial velocity at top of reactor, U f, m/s Solids circulation flux at top of reactor, G sf, kg/m2 s Density ratio, q p/q g Froude number, Fr Archimedes number, Ar Solids-to-gas mass flow ratio, G sf /(q gU f) Inventory of solids, H/D
Commercial Syncrude units, base conditions
UBC, base conditions with FCC
Vaporized hydrocarbons and steam 510–540 360 2.28 2.5105 Fluid Coke 1600 145 A 0.8
Air
25 200 2.34 1.8105 FCC 1700 99 A 0.74
20.3
18.6
701 485 174 10.7
714 567 119 10.5
2.18
2.15
a
As calculated at the operating pressure and temperature from the equations given by Grace [17].
ing in diameter from 305 mm at the bottom to 483 mm at the top. Despite the tapered expansion, the superficial gas velocity increases up the feed section, due to stagewise addition of gas through a series of injection rings. The geometry of the reactor section is identical to that of the commercial unit, but scaled down by a factor of ~20. In the commercial units, a mixture of bitumen and steam is injected into the reactor section through six rings of nozzles. The jet geometry is assumed to be composed of a cone and a hemisphere attached to the base of the cone, with the jet half-angle assumed constant and equal to 6.258 [9]. The jet penetration depth, L, defined as the horizontal distance from the tip of the nozzle to the mean position of the end of the jet region, can be estimated from the semiempirical correlation of Merry [9] 0:4 0:2 0:2 qg dp L q0 u20 þ 4:5 ¼ 5:25 d0 ð1 eÞqp gdp qp d0 ð2Þ This correlation has previously been shown to provide good estimation of the horizontal jet penetration [1,6,10,11]. The nozzle tip velocity in the cold model was set to maintain the same ratio of jet penetration to column radius as for the commercial units, thereby matching the dimensionless radial positions of the ends of the jets. The reactor section contains six rows of injection nozzles at different levels. The volumetric flowrate through each ring is measured by an orifice flowmeter and adjusted by a globe valve. Each feed ring contains numerous radially orientated nozzles, which
distribute the air uniformly among the nozzles in that ring. Polyethylene tubes of 12.7 mm OD of the same length for each ring connect each nozzle to a manifold. To control each nozzle individually and to prevent the lines from plugging, ball valves are installed upstream of each nozzle. An additional ring of nozzles is installed between the reactor and stripping section to simulate the high-velocity steam attrition jets used to maintain a stable particle size distribution in the commercial unit. Instrumentation ports are provided at nine axial locations in the reactor section. Pressure taps are located at each instrumentation port level, with differential pressure transducers (Omega, PX-164) monitoring the pressure gradient distribution along the height of the reactor section. Three levels of nozzles in the lower region supply the commercial strippers with stripping steam and provide the fluidization gas needed in this section. In the cold model, compressed air replaces steam in the stripping section. After descending through the reactor and stripper sections, particles leave the stripper into a standpipe and then travel through a pinch valve and U-bend. The solids are then transported through an external vertical riser (see Fig. 1) and returned to the freeboard of the reactor section. The solids circulation rate is controlled by the pinch valve and monitored by recording the pressure drop across a venturi constriction near the exit of the riser. The pressure drop vs. solids circulation rate was calibrated by determining the particle velocity and the voidage in the vertical standpipe by optical fibre probes for the same riser gas flow rate as in the hydrodynamic tests. A particle–gas separator is installed at the exit of the riser. Further separation of exhaust air is accomplished by six parallel cyclones. The solids captured by the primary cyclones return to the unit, while those not caught pass to a secondary cyclone and finally to two filter bag-houses in parallel. A pinch valve is installed in the exhaust pipe downstream of the primary cyclones to control the operating pressure in order to achieve the desired solidto-gas density ratio. 2.1. Fibre optical voidage measurement system A reflective-type optical fibre probe, made from two interspersed bundles of quartz fibres encased in a 0.5 m long, 4.8 mm OD stainless steel tube with an active sensor area of 22 mm, was used to measure the local voidage in the reactor section. The probe emits light into a fluidized bed of FCC particles, and the local concentration of the particles at each radial position is determined from the intensity of the reflected light at a sampling frequency of 100 Hz for periods of 80 s. The concentration probe was calibrated using mixtures of FCC and fluid coke particles for known volume fractions of FCC. Fluid coke particles have poor light reflection and consequently contribute very little to the light reflection from FCC–coke mixtures. As fluid cokers are operated in the bubbling and turbulent fluidization flow regimes, the
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1
calibration measurements had to cover a wide range of FCC concentrations right up to the loose-packed-bed value. The calibration curve was also confirmed by comparing the integrated cross-sectional average value with that calculated from the pressure drop across a measurement interval in a column with vertical walls.
0.9 0.8 0.7
2.2. Fibre optical particle velocity measurement system An optical probe with three fibres was used to measure the velocity of particles in the reactor section. Details of this probe have been presented elsewhere [12,13]. The central fibre was used to illuminate moving particles, while the other two fibres capture light reflected by passing particles. The local velocity component is determined from the transit time and the effective distance between the receiving fibres. The transit time was calculated by cross-correlating the signals from the two channels, while the effective distance was calibrated with the aid of a disc, onto which particles were glued, rotating at a known speed. The effective distance of the probe was 0.23 mm, about 2.3 times the average FCC particle size. For all measurements in the wall region (r/RN0.7), data were obtained at frequencies of 7.8 and 15.6 kHz for sampling periods of 40 to 80 s. For the core region (r/Rb0.7), the frequencies were 31.3 and 62.5 kHz and the sampling period was 10 to 20 s. Fewer than 5% of the data had to be rejected due to poor correlation between the channels.
3. Results and discussion 3.1. Axial pressure gradient Pressure fluctuations provide one of the easiest and least costly means of studying the hydrodynamics of fluidized beds [14,15]. Differential pressure measurements, with the separation distance between ports of the order of a few centimeters, are reported to reflect the local flow behaviour. Pressure gradients in the reactor section were measured at different heights. Dynamic differential pressure fluctuations were recorded at a sampling frequency of 100 Hz for durations of 85 s. Time-averaged pressure gradients in the reactor section under base conditions appear in Fig. 3. As the gas velocity increases with increasing height, the pressure gradient decreases, as expected. Note that the cross-sectional average voidage cannot be simply calculated from the pressure gradient because of the taper (non-vertical walls). Power spectral diagrams of differential pressure for four axial positions in the reactor section are provided in Fig. 4. In the lower part of the reactor section, a strong peak is found in the range of 0 to 5 Hz. This peak shifts to lower frequencies in the upper part because of the coalescence of voids and the increased superficial gas velocity (due to air injection from the feed nozzles). A similar trend was
Z* [-]
0.6 0.5 0.4 0.3 0.2 0.1 0 0
2
4
6
8
10
dP/dZ [kPa/m] Fig. 3. Axial time-mean pressure gradient profiles in reactor section. Base conditions, U f =0.74 m/s, U s=0.25 m/s, G sf=18.6 kg/m2 s.
observed by Ellis et al. [15] based on FFT analysis of differential pressure fluctuation signals at several axial positions in fluidized beds of different diameter. 3.2. Voidage Distributions Earlier work in the semi-cylindrical Plexiglas column referred to above indicates that the reactor section operated near the boundary between bubbling and turbulent fluidization, with a relatively dense region of downward flow along the wall coupled with accelerated upward flow at higher voidage in the interior of the column [1,6]. The radial non-uniformity is more pronounced in the model fluid coker when air is injected radially from the nozzles. Fig. 5 shows a typical radial profile of time-mean voidage for given radial positions. The time-mean voidage increases gradually from the wall to the axis of the column. The inserts show the nature of the fluctuations for four radial positions. Significant peaks in voidage in these inserts correspond to voids of low solids concentration. The frequency of these voids increases from the wall to the centre. Fig. 6 shows the local voidage in the reactor section as a function of height and radius under base operating conditions. It can be seen that there is no distinct interface between a dilute core and a dense annular region. At the lower end near the stripper (port X8), the radial profile of the voidage is relatively uniform because of baffles (sheds) in the stripper section. The void size is then restricted by the internal baffles. As the gas rises, voids coalesce with feed jets and their size increases dramatically in the core region. The voidage increases gradually from the outer annular zone to the core. The results also show slight changes with height caused by the axial variation of superficial gas velocity, causing the bed to be denser in the lower part and more dilute in the upper part of the reactor section. This is
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Fig. 5. Typical radial profile of time-mean voidage and local voidage fluctuations measured by optical fibre probe at measurement port Y4 (between feed rings 3 and 4) for U g=0.54 m/s and G s=24.4 kg/m2 s.
Fig. 4. Power spectrum diagrams of local differential pressure fluctuations in reactor section. Base conditions; U f =0.74 m/s, U s=0.25 m/s, G sf=18.6 kg/m2 s.
consistent with the pressure gradient distribution shown in Fig. 3. The effect of changing the solids circulation flux was examined by varying G sf from 0 to 28.7 kg/m2 s. The radial voidage profiles were measured at one level in the reactor section, as shown in Fig. 7. It can be seen that there was little
Fig. 6. Voidage distribution in reactor section. Base conditions, U f =0.74 m/s, U s=0.25 m/s, G sf=18.6 kg/m2 s. Feed jet penetration calculated from Eq. (2).
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Fig. 7. Radial voidage profiles for different solids circulation fluxes at U g=0.42 m/s.
change in the voidage profiles with changing solids circulation flux, G sf. The effect of the superficial gas velocity on the radial voidage distribution is discussed below. 3.3. Solids velocity and flow structure Matsen [2] reported that solids mixing in large units like fluid cokers is governed by intense gulf-stream circulation patterns, up the centre and down at the walls. The gulf streaming concentrates the gas bubbles or voids in the centre of the column and enhances their velocities. Knapper et al. [1] applied a core-annulus model to the reactor section of the coker with a dilute, upward-flowing stream of gas in the central (core) region of the reactor
assumed to be surrounded by a dense, downward-flowing (annular) region of particles. The boundary of the coreannular region can be defined as the position where the net local vertical solids flux or the time-mean vertical component of particle velocity is zero. It must be noted [16] that the time-mean flux is in general not equal to the product of the time-mean particle velocity and time-mean bulk density, i.e. Z P P 1 T P Gs ¼ qP VP ðt Þ½1 eðt Þdt pqP VP 1 e ð3Þ T 0 This arises because the fluctuations in voidage or solids concentration tend to be strongly correlated with fluctuations in particle velocity. Hence, basing the core/annulus
Fig. 8. Typical radial profile of time-mean particle velocity and probability distributions for three radial locations. Measurement port: Y4 (between feed rings 3 and 4), U g=0.48 m/s, G s=24.4 kg/m2 s.
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boundary on where VP passes through zero gives Pa different boundary position than basing it on where Gs passes through zero [16]. Because of the nature of the correlation between voidage and particle velocity fluctuations, the wall layer thickness based on particle velocity measurements is usually expected to be smaller than that determined from solid flux measurements [16]. In this work, the boundary is based on the radial position at which the time-mean vertical velocity component passes through zero. A typical radial profile of time-mean particle velocity for a given set of operating conditions is shown in Fig. 8. The probability distributions of particle velocity for three radial locations are also shown in this figure. They were calculated by counting the relative number of times the
measured particle velocity falls in 0.2 m/s intervals covering the range from 2.0 to 5.0 m/s. The direction of particle velocity is mainly upward in the core region, and mainly downward in the annular region, with a wide spread of both upward and downward instantaneous values in the transition zone between the core and annular regions. The downward flow fraction, i.e. the fraction of the measured particle velocities which are downward, is also plotted in this figure. The positive velocities in the core region indicate upward flow of solids. With increasing r/R, the time-mean particle velocity gradually decreases and finally becomes negative. An overall flow map for the reactor section is shown in Fig. 9. At the top of the stripper section, i.e. port X8, it
Fig. 9. Overall distribution of time-average vertical velocity component in reactor section. Base conditions, U f =0.74 m/s, U s=0.25 m/s, G sf=18.6 kg/m2 s. Feed jet penetration calculated from Eq. (2).
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133
Fig. 10. Boundary between upward and downward velocity in reactor section compared with penetration of feed nozzles and jets into the column.
is unclear whether or not there is a wall zone with net downward velocity, although there is definitely a somewhat higher upward velocity in the central region. The wall region at this level is subject to flow reversal, with solids traveling intermittently upward and downward, resulting in extensive turbulence. The gas leaving the stripper section is well dispersed throughout the crosssection by the baffles, with few large bubbles or voids. As shown in Fig. 10, the ratio of the area of the downwardparticle-velocity region to the total cross-sectional area increases with height in the reactor section, reaching a maximum about half-way up the column and decreasing towards the top, probably due to turbulence caused by higher superficial gas velocity there. The average dimensionless radial position (r/R) of this boundary is 0.77, suggesting that the downward-particle-velocity region occupies about 40% of the cross-sectional area on
Fig. 12. Effect of external solids circulation flux on particle velocity distribution measured at the port between feed rings 3 and 4 with U g=0.48 m/s at this level.
average. The radial positions of the nozzle tips and the jet penetration, calculated from Merry [9], are also shown in Fig. 10. The feed jets are seen to penetrate through the transition zone between the core and annular regions. Particles descending in the annular region are entrained by the jets into the upward flowing core region. While ascending, some particles transfer from the core to the wall region where they mainly descend along the column wall. The effect of nozzle insertion distance was investigated by retracting all feed nozzles from an insertion distance of 29 mm to be flush with the wall. Results from measurement port Y4 (between feed rings 3 and 4) are shown in Fig. 11. It can be seen that the annular region where particles descend became narrower, while the radial profile of time-mean particle velocity became less non-uniform over the cross-section. The effect of the solids circulation rate on the time-mean particle velocity profile was investigated by varying G s from 0 to 30.3 kg/m2 s at a level between feed rings 3 and 4. Results are shown in Fig. 12. The cross-sectional average
Table 2 Superficial velocities in reactor section
Fig. 11. Effect of nozzle insertion distance for G s=24.4 kg/m2 s at measurement port Y4 between feed rings 3 and 4 with U g=0.48 m/s at this level.
Commercial baseline
Lower feed rates
Higher feed rates
U f at top of reactor
0.74 m/s
Measurement port Y2 between feed rings 1 and 2 Measurement port Y4 between feed rings 3 and 4
0.63 m/s
0.52 m/s (30%) 0.45 m/s (29%)
0.98 m/s (+32%) 0.81 m/s (+29%)
0.38 m/s (21%)
0.59 m/s (+23%)
0.48 m/s
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Fig. 13. Impact of feed rates through nozzles on radial voidage profiles in reactor. For superficial gas velocities see Table 2.
downward particle velocity caused by the external circulation through the riser is 0.05 m/s for the base conditions where G s=24.4 kg/m2 s at this level. Increasing the solids circulation flux is seen to have caused a slight widening of the annular region of the bed in which the net movement is downward. 3.4. Impact of lower and higher feed rate The impact of the superficial gas velocity on the solid flow was investigated by varying the gas flow through each of the feed jets by ~F40% compared with the base conditions. The superficial gas velocities at the measurement ports are compared in Table 2. Note that the magnitudes of the total changes are less than F40% because
the flow from the stripper remained unchanged. Radial profiles of time-mean voidage determined from ports Y2 (between feed rings 1 and 2) and Y4 (between feed rings 3 and 4) are shown in Fig. 13. At both levels and for all radial positions, the bed became more dilute at higher nozzle feed rates and denser at lower feed rates. The shapes of the profile are similar, with a dense annular region surrounding a more dilute core. The radial profiles of standard deviation of local voidage in Fig. 13 indicate similar variation across the cross-section. Time-mean particle velocity profiles are shown in Fig. 14. The boundary separating upward and downward particle mean motion did not change significantly with the total gas flow rate. However, the particle velocity in the core region did depend on the gas flow, especially in the upper region.
Fig. 14. Impact of feed rates through nozzles on solids flow in reactor. For superficial gas velocities see Table 2.
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The jet penetrations calculated from Merry’s equation are also indicated in Fig. 14. In the upper part of the column, as shown in Fig. 14(a), the particle velocity in the central region decreased dramatically and the radial distribution became flatter when the system was operated at lower gas flow rates, presumably because of reduced jet penetration. The internal circulation between the core and annular regions decreased with decreasing superficial gas velocity.
Uf Ug Us VP Z Z*
4. Conclusions Axial and radial distributions of voidage in the reactor section of a geometrically and dynamically scaled fluid coker cold model were measured by an optical fibre probe. The results show a dense annular region surrounding a more dilute core region. The voidage profiles indicate that the voidage increases smoothly from the wall to the axis of the column, rather than showing a sharp transition between annular and core regions. By measuring the particle velocity in the reactor section with another optical fibre probe, it was confirmed that particles tend to rise in the central core, while descending in the outer annular region. The relative radial position where the time-mean vertical velocity component is zero changes with height. For the tapered column (geometrically similar to industrial reactors) used in this work, the downward-flow annular region occupied ~40% of the overall cross-sectional area. Particles descending in the annular region were entrained by the feed jet into the upward-flowing core region. The boundary between upward and downward velocity regions did not change appreciably when the jet gas flow rates were varied by ~F40% from the base conditions. However, the time-mean particle velocity and internal circulation between the core and annular region did change significantly over this range.
Nomenclature P3 q ðq q Þgdp Ar Archimedes number, ¼ g P l2g D diameter at top of reactor section, m d inner diameter of nozzle, m 0 P dp mean particle diameter, m F cumulative mass fraction under U2 Fr Froude number, ¼ g dPf p g acceleration due to gravity, 9.81 m/s2 Gs local net solids circulation flux corrected for crosssectional area, kg/m2 s G sf net solids circulation flux based on cross-sectional area at top of reactor, kg/m2 s H expended bed depth in reactor section, m L horizontal jet penetration, m t time, s T total sampling duration, s u0 feed nozzle jet velocity, m/s
Greek e l qo qg qp
135
superficial gas velocity at dense phase bed surface, m/s local superficial gas velocity corrected for gas flowrate and cross-sectional area, m/s superficial gas velocity at top of stripper section, m/s vertical particle velocity, m/s height coordinate measured from top of highest stripper shed, m dimensionless height coordinate (height coordinate divided by total height of reactor section) letters local voidage gas viscosity, Pa s gas density at nozzle, kg/m3 gas density in fluidized bed, kg/m3 particle density, kg/m3
Acknowledgement The authors are grateful to Syncrude Canada for sponsoring this work and for permission to publish the results. We also wish to thank D. Famulak, J. Xu and W. Wei for their assistance with experiments and Chevron Canada for providing the FCC particles. References [1] B. Knapper, F. Berruti, J.R. Grace, H.T. Bi, C.J. Lim, Hydrodynamic characterization of fluid bed cokers, in: J.R. Grace, J. Zhu, H. de Lasa (Eds.), Circulating Fluidized Bed Technology VII, C.S.Ch.E., Ottawa, 2002, pp. 263 – 270. [2] J.M. Matsen, Scale-up of fluidized bed process: principle and practice, Powder Technol. 88 (1996) 237 – 244. [3] M.R. Gray, T. Le, W.C. McCaffrey, F. Berruti, S. Soundararajan, E. Chan, I. Huq, Coupling of mass transfer and reaction in coking of thin films of Athabasca vacuum residue, Ind. Eng. Chem. Res. 40 (2001) 3317 – 3324. [4] X.T. Bi, H.P. Cui, J.R. Grace, A. Kern, C.J. Lim, D. Rusnell, X.Q. Song, C. McKnight, Flooding of gas–solids counter-current flow in fluidized beds, Ind. Eng. Chem. Res. 43 (2004) 5611 – 5619. [5] X.T. Bi, J.R. Grace, C.J. Lim, C. McKnight, Hydrodynamics of the stripper section of fluid cokers, 2004, in press. [6] B.A. Knapper, Experimental studies on the hydrodynamics of fluid bed cokers, MSc thesis, University of Saskatchewan, Saskatoon, Canada, 2000. [7] L.R. Glicksman, M.R. Hyre, P.A. Farrell, Dynamic similarity in fluidization, Int. J. Multiph. Flow 20S (1994) 331 – 386. [8] P.A. Farrell, M.R. Hyre, L.R. Glicksman, Importance of the solid-togas density ratio for scaling fluidized bed hydrodynamics, in: L.S. Fan, T.M. Knowlton (Eds.), Fluidization IX, Engineering Foundation, New York, 1998, pp. 85 – 92. [9] J.M.D. Merry, Penetration of a horizontal gas jet into a fluidized bed, Trans. Inst. Chem. Eng. 49 (1971) 189 – 195. [10] J. Copan, Macroscopic modeling of a fluid bed coker and experimental studies of one- and two-phase feed jets, MSc thesis, University of Saskatchewan, Saskatoon, Canada, 1999. [11] A. Donald, H.T. Bi, J.R. Grace, C.J. Lim, Penetration of single and multiple horizontal jets into fluidized beds, Fluidization XI, Ischia (2004) 171 – 178.
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[12] J. Liu, J.R. Grace, X.T. Bi, Novel multifunctional optical-fiber probe: I. Development and validation, AIChE J. 49 (2003) 1405 – 1420. [13] J. Liu, J.R. Grace, X.T. Bi, Novel multifunctional optical-fiber probe: II. High-density CFB measurements, AIChE J. 49 (2003) 1421 – 1432. [14] H.T. Bi, N. Ellis, I.A. Abba, J.R. Grace, A state-of-the-art review of gas–solid turbulent fluidization, Chem. Eng. Sci. 55 (2000) 4789 – 4825. [15] N. Ellis, C.J. Lim, J.R. Grace, H.T. Bi, K.S. Lim, Frequency analysis of pressure fluctuations in turbulent fluidized beds, in: J.R. Grace,
J. Zhu, H. de Lasa (Eds.), Circulating Fluidized Bed Technology VII, C.S.Ch.E., Ottawa, 2002, pp. 287 – 294. [16] H.T. Bi, J. Zhou, S.Z. Qin, J.R. Grace, Annular wall layer thickness in circulating fluidized bed risers, Can. J. Chem. Eng. 74 (1996) 811 – 814. [17] J.R. Grace, Contacting modes and behaviour classification of gas– solid and other two-phase suspensions, Can. J. Chem. Eng. 64 (1986) 353 – 363.
Hydrodynamics of the reactor section in fluid cokers Xuqi Songa, Hsiaotao Bia, C. Jim Lima, John R. Gracea,*, Edward Chanb, Brian Knapperb, Craig McKnightb a
Fluidization Research Centre, Department of Chemical and Biological Engineering, University of British Columbia, 2216 Main Mall, Vancouver, British Columbia, Canada V6T 1Z4 b Syncrude Research Centre, 9421-17th Avenue, Edmonton, Alberta, Canada T6N 1H4 Received 6 January 2004; received in revised form 18 August 2004; accepted 17 September 2004 Available online 6 November 2004
Abstract The hydrodynamics of fluid cokers were studied in a pressurized fully cylindrical cold model scaled by matching key dimensionless groups so that the reactor sections of the commercial and laboratory units were geometrically and dynamically similar. Voidage distributions and solids flow structure were investigated using optical fibre probes in a fluidized bed of FCC. The results show a relatively dense annular region, in which the time-average flow is downwards, surrounding a more dilute, upward flowing core region. However, the radial profiles of voidage indicate that the voidage increases gradually from the wall to the axis of the column, rather than showing a sharp transition between an annular and core region. Particles in the descending outer region are entrained by the feed jets into the core region. The radial position of the boundary between upward and downward net solids flow did not change significantly when the superficial gas velocity was varied. D 2004 Elsevier B.V. All rights reserved. Keywords: Fluidization; Fluid coker; Hydrodynamics; Voidage
1. Introduction Fluid coking is a non-catalytic pyrolysis process used to convert high-molecular-weight hydrocarbons to lower boiling point hydrocarbons, gases and coke. The thermal cracking process is carried out in large fluidized beds, where hot coke particles, introduced from above, carry the heat required for the endothermic reactions and collect solid byproducts on their surfaces. The hydrocarbon feed is injected through several rows of horizontal nozzles and the vapor products tend to migrate up the centre of the bed, surrounded by dense, descending particles [1]. Particles descend through the reactor and then through a stripper, in which a countercurrent flow of steam strips liquid hydrocarbon from the particle surface. The coke particles are then circulated into a fluidized bed burner for re-heating
* Corresponding author. Tel.: +1 604 822 3121; fax: +1 604 822 6003. E-mail address: [email protected] (J.R. Grace). 0032-5910/$ - see front matter D 2004 Elsevier B.V. All rights reserved. doi:10.1016/j.powtec.2004.09.033
prior to re-entering the fluid coker at an increased temperature [2]. After being steam-atomized and injected through feed nozzles, the heavy hydrocarbon feed remains as a thin liquid film on coke particles within the reactor until the feed has fully reacted to form vapour products or coke byproduct. Thin films of feed on coke provide maximum liquid yield [3]. The distribution of liquid feed over the coke solids is strongly dependent on the mode in which the spray from the feed nozzles interacts with the surrounding particles. In order to optimize and improve the performance, it is important to understand the radial distribution of solids concentration and the particle flow structure within the reactor. To improve the understanding of the hydrodynamics in the coker, two cold flow models (a semi-circular Plexiglas column and a fully cylindrical steel column) were designed, constructed and commissioned at the University of British Columbia. The semi-circular cold model has been used primarily to study the hydrodynamics of the stripper section
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to guide the operation and improvement in baffle configuration and solids circulation. Several operational problems, such as fouling and flooding have been investigated in this equipment [4,5]. Measurements of key characteristic parameters in the reactor section were also made for various operating conditions in the semicircular cold model [1,6]. However, the semi-cylindrical column was designed to provide similarity with commercial cokers in the stripper section, not the reactor section. Moreover, being semicylindrical, the column introduces wall and corner effects which make it difficult to apply the results to the fully cylindrical commercial cokers. Indeed the flat face of the semi-cylindrical column was found to have a significant effect on the cross-sectional distribution of voidage and axial solids flow patterns [6]. A pressurized cold model facility was therefore designed, constructed and commissioned to examine the hydrodynamic behaviour in the reactor section of a column whose cross-section was fully circular.
2. Experimental An overall schematic of the fully cylindrical cold-model facility is shown in Fig. 1. All components were fabricated of steel. These include a reactor section and a stripper section on the coker side, a riser for external solids circulation and a collection system for entrained particles. The equipment was designed so that its reactor section would be both geometri-
Fig. 1. Schematic of UBC pressurized fully cylindrical cold model of Syncrude Cokers.
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Fig. 2. Cumulative mass fractions for FCC and Fluid Coke particles.
cally and dynamically similar to those of two commercial fluid cokers operated by Syncrude Canada Limited in Fort McMurray, Alberta, Canada. Scaling, based on dimensionless analysis [7] was achieved by matching the important dimensionless hydrodynamic parameters Ar; Fr;
qp Gs ; ; PSD; H=D qg qg Uf
ð1Þ
Due to the large geometric scale-down factor and the fact that the particles used in the commercial cokers are already quite small, it was not possible to keep the ratio of D/d p constant. However, D/d pNN1 in both cases, so that D/d p matching can be relaxed, so long as all gaps through which the particles flow in the reactor and stripper sections are wider than 50d p~. Above the bed surface the cold model column expands beyond the geometric similarity dimensions in order to assist in disengaging particles in the freeboard zone. Plots of the cumulative mass fraction for the FCC and fluid coke particles are shown in Fig. 2. FCC particles were chosen for this study because optical fibre probes can be used to investigate hydrodynamics of these nearly white particles in the cold model, whereas this is impossible with the black coke. Other properties appear in Table 1. The dimensionless PSDs are quite similar for the FCC and coke particles, as required by the scaling criteria. The results of scaling calculations based on the superficial gas velocity at the surface of the dense bed in the reactor section are summarized in Table 1. The bbase conditionsQ referred to hereafter are the operating conditions (with all 6 feed rings operating, U s=0.25 m/s; U f =0.74 m/s; G sf=18.6 kg/m2 s) having dynamic similitude with the commercial units under normal operation. Experiments by Farrell et al. [8] showed that the solid-togas density ratio must be matched to ensure accurate scaling of the hydrodynamics in bubbling or turbulent fluidized beds. To maintain the same density ratio, q p/q g, as in the commercial units, the cold model was operated at a slightly elevated pressure, 200 kPa, at the bottom of the reactor section. This section is axisymmetric and tapered, expand-
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Table 1 Summary of dimensional similitude based on reactor section
Gas
Temperature, 8C Pressure, kPa Gas density, kg/m3 Gas viscosity, Pa s Type of particles Particle density, kg/m3 Mean particle diameter, Am Geldart powder groupa Superficial velocity at top of reactor, U f, m/s Solids circulation flux at top of reactor, G sf, kg/m2 s Density ratio, q p/q g Froude number, Fr Archimedes number, Ar Solids-to-gas mass flow ratio, G sf /(q gU f) Inventory of solids, H/D
Commercial Syncrude units, base conditions
UBC, base conditions with FCC
Vaporized hydrocarbons and steam 510–540 360 2.28 2.5105 Fluid Coke 1600 145 A 0.8
Air
25 200 2.34 1.8105 FCC 1700 99 A 0.74
20.3
18.6
701 485 174 10.7
714 567 119 10.5
2.18
2.15
a
As calculated at the operating pressure and temperature from the equations given by Grace [17].
ing in diameter from 305 mm at the bottom to 483 mm at the top. Despite the tapered expansion, the superficial gas velocity increases up the feed section, due to stagewise addition of gas through a series of injection rings. The geometry of the reactor section is identical to that of the commercial unit, but scaled down by a factor of ~20. In the commercial units, a mixture of bitumen and steam is injected into the reactor section through six rings of nozzles. The jet geometry is assumed to be composed of a cone and a hemisphere attached to the base of the cone, with the jet half-angle assumed constant and equal to 6.258 [9]. The jet penetration depth, L, defined as the horizontal distance from the tip of the nozzle to the mean position of the end of the jet region, can be estimated from the semiempirical correlation of Merry [9] 0:4 0:2 0:2 qg dp L q0 u20 þ 4:5 ¼ 5:25 d0 ð1 eÞqp gdp qp d0 ð2Þ This correlation has previously been shown to provide good estimation of the horizontal jet penetration [1,6,10,11]. The nozzle tip velocity in the cold model was set to maintain the same ratio of jet penetration to column radius as for the commercial units, thereby matching the dimensionless radial positions of the ends of the jets. The reactor section contains six rows of injection nozzles at different levels. The volumetric flowrate through each ring is measured by an orifice flowmeter and adjusted by a globe valve. Each feed ring contains numerous radially orientated nozzles, which
distribute the air uniformly among the nozzles in that ring. Polyethylene tubes of 12.7 mm OD of the same length for each ring connect each nozzle to a manifold. To control each nozzle individually and to prevent the lines from plugging, ball valves are installed upstream of each nozzle. An additional ring of nozzles is installed between the reactor and stripping section to simulate the high-velocity steam attrition jets used to maintain a stable particle size distribution in the commercial unit. Instrumentation ports are provided at nine axial locations in the reactor section. Pressure taps are located at each instrumentation port level, with differential pressure transducers (Omega, PX-164) monitoring the pressure gradient distribution along the height of the reactor section. Three levels of nozzles in the lower region supply the commercial strippers with stripping steam and provide the fluidization gas needed in this section. In the cold model, compressed air replaces steam in the stripping section. After descending through the reactor and stripper sections, particles leave the stripper into a standpipe and then travel through a pinch valve and U-bend. The solids are then transported through an external vertical riser (see Fig. 1) and returned to the freeboard of the reactor section. The solids circulation rate is controlled by the pinch valve and monitored by recording the pressure drop across a venturi constriction near the exit of the riser. The pressure drop vs. solids circulation rate was calibrated by determining the particle velocity and the voidage in the vertical standpipe by optical fibre probes for the same riser gas flow rate as in the hydrodynamic tests. A particle–gas separator is installed at the exit of the riser. Further separation of exhaust air is accomplished by six parallel cyclones. The solids captured by the primary cyclones return to the unit, while those not caught pass to a secondary cyclone and finally to two filter bag-houses in parallel. A pinch valve is installed in the exhaust pipe downstream of the primary cyclones to control the operating pressure in order to achieve the desired solidto-gas density ratio. 2.1. Fibre optical voidage measurement system A reflective-type optical fibre probe, made from two interspersed bundles of quartz fibres encased in a 0.5 m long, 4.8 mm OD stainless steel tube with an active sensor area of 22 mm, was used to measure the local voidage in the reactor section. The probe emits light into a fluidized bed of FCC particles, and the local concentration of the particles at each radial position is determined from the intensity of the reflected light at a sampling frequency of 100 Hz for periods of 80 s. The concentration probe was calibrated using mixtures of FCC and fluid coke particles for known volume fractions of FCC. Fluid coke particles have poor light reflection and consequently contribute very little to the light reflection from FCC–coke mixtures. As fluid cokers are operated in the bubbling and turbulent fluidization flow regimes, the
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1
calibration measurements had to cover a wide range of FCC concentrations right up to the loose-packed-bed value. The calibration curve was also confirmed by comparing the integrated cross-sectional average value with that calculated from the pressure drop across a measurement interval in a column with vertical walls.
0.9 0.8 0.7
2.2. Fibre optical particle velocity measurement system An optical probe with three fibres was used to measure the velocity of particles in the reactor section. Details of this probe have been presented elsewhere [12,13]. The central fibre was used to illuminate moving particles, while the other two fibres capture light reflected by passing particles. The local velocity component is determined from the transit time and the effective distance between the receiving fibres. The transit time was calculated by cross-correlating the signals from the two channels, while the effective distance was calibrated with the aid of a disc, onto which particles were glued, rotating at a known speed. The effective distance of the probe was 0.23 mm, about 2.3 times the average FCC particle size. For all measurements in the wall region (r/RN0.7), data were obtained at frequencies of 7.8 and 15.6 kHz for sampling periods of 40 to 80 s. For the core region (r/Rb0.7), the frequencies were 31.3 and 62.5 kHz and the sampling period was 10 to 20 s. Fewer than 5% of the data had to be rejected due to poor correlation between the channels.
3. Results and discussion 3.1. Axial pressure gradient Pressure fluctuations provide one of the easiest and least costly means of studying the hydrodynamics of fluidized beds [14,15]. Differential pressure measurements, with the separation distance between ports of the order of a few centimeters, are reported to reflect the local flow behaviour. Pressure gradients in the reactor section were measured at different heights. Dynamic differential pressure fluctuations were recorded at a sampling frequency of 100 Hz for durations of 85 s. Time-averaged pressure gradients in the reactor section under base conditions appear in Fig. 3. As the gas velocity increases with increasing height, the pressure gradient decreases, as expected. Note that the cross-sectional average voidage cannot be simply calculated from the pressure gradient because of the taper (non-vertical walls). Power spectral diagrams of differential pressure for four axial positions in the reactor section are provided in Fig. 4. In the lower part of the reactor section, a strong peak is found in the range of 0 to 5 Hz. This peak shifts to lower frequencies in the upper part because of the coalescence of voids and the increased superficial gas velocity (due to air injection from the feed nozzles). A similar trend was
Z* [-]
0.6 0.5 0.4 0.3 0.2 0.1 0 0
2
4
6
8
10
dP/dZ [kPa/m] Fig. 3. Axial time-mean pressure gradient profiles in reactor section. Base conditions, U f =0.74 m/s, U s=0.25 m/s, G sf=18.6 kg/m2 s.
observed by Ellis et al. [15] based on FFT analysis of differential pressure fluctuation signals at several axial positions in fluidized beds of different diameter. 3.2. Voidage Distributions Earlier work in the semi-cylindrical Plexiglas column referred to above indicates that the reactor section operated near the boundary between bubbling and turbulent fluidization, with a relatively dense region of downward flow along the wall coupled with accelerated upward flow at higher voidage in the interior of the column [1,6]. The radial non-uniformity is more pronounced in the model fluid coker when air is injected radially from the nozzles. Fig. 5 shows a typical radial profile of time-mean voidage for given radial positions. The time-mean voidage increases gradually from the wall to the axis of the column. The inserts show the nature of the fluctuations for four radial positions. Significant peaks in voidage in these inserts correspond to voids of low solids concentration. The frequency of these voids increases from the wall to the centre. Fig. 6 shows the local voidage in the reactor section as a function of height and radius under base operating conditions. It can be seen that there is no distinct interface between a dilute core and a dense annular region. At the lower end near the stripper (port X8), the radial profile of the voidage is relatively uniform because of baffles (sheds) in the stripper section. The void size is then restricted by the internal baffles. As the gas rises, voids coalesce with feed jets and their size increases dramatically in the core region. The voidage increases gradually from the outer annular zone to the core. The results also show slight changes with height caused by the axial variation of superficial gas velocity, causing the bed to be denser in the lower part and more dilute in the upper part of the reactor section. This is
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Fig. 5. Typical radial profile of time-mean voidage and local voidage fluctuations measured by optical fibre probe at measurement port Y4 (between feed rings 3 and 4) for U g=0.54 m/s and G s=24.4 kg/m2 s.
Fig. 4. Power spectrum diagrams of local differential pressure fluctuations in reactor section. Base conditions; U f =0.74 m/s, U s=0.25 m/s, G sf=18.6 kg/m2 s.
consistent with the pressure gradient distribution shown in Fig. 3. The effect of changing the solids circulation flux was examined by varying G sf from 0 to 28.7 kg/m2 s. The radial voidage profiles were measured at one level in the reactor section, as shown in Fig. 7. It can be seen that there was little
Fig. 6. Voidage distribution in reactor section. Base conditions, U f =0.74 m/s, U s=0.25 m/s, G sf=18.6 kg/m2 s. Feed jet penetration calculated from Eq. (2).
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Fig. 7. Radial voidage profiles for different solids circulation fluxes at U g=0.42 m/s.
change in the voidage profiles with changing solids circulation flux, G sf. The effect of the superficial gas velocity on the radial voidage distribution is discussed below. 3.3. Solids velocity and flow structure Matsen [2] reported that solids mixing in large units like fluid cokers is governed by intense gulf-stream circulation patterns, up the centre and down at the walls. The gulf streaming concentrates the gas bubbles or voids in the centre of the column and enhances their velocities. Knapper et al. [1] applied a core-annulus model to the reactor section of the coker with a dilute, upward-flowing stream of gas in the central (core) region of the reactor
assumed to be surrounded by a dense, downward-flowing (annular) region of particles. The boundary of the coreannular region can be defined as the position where the net local vertical solids flux or the time-mean vertical component of particle velocity is zero. It must be noted [16] that the time-mean flux is in general not equal to the product of the time-mean particle velocity and time-mean bulk density, i.e. Z P P 1 T P Gs ¼ qP VP ðt Þ½1 eðt Þdt pqP VP 1 e ð3Þ T 0 This arises because the fluctuations in voidage or solids concentration tend to be strongly correlated with fluctuations in particle velocity. Hence, basing the core/annulus
Fig. 8. Typical radial profile of time-mean particle velocity and probability distributions for three radial locations. Measurement port: Y4 (between feed rings 3 and 4), U g=0.48 m/s, G s=24.4 kg/m2 s.
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boundary on where VP passes through zero gives Pa different boundary position than basing it on where Gs passes through zero [16]. Because of the nature of the correlation between voidage and particle velocity fluctuations, the wall layer thickness based on particle velocity measurements is usually expected to be smaller than that determined from solid flux measurements [16]. In this work, the boundary is based on the radial position at which the time-mean vertical velocity component passes through zero. A typical radial profile of time-mean particle velocity for a given set of operating conditions is shown in Fig. 8. The probability distributions of particle velocity for three radial locations are also shown in this figure. They were calculated by counting the relative number of times the
measured particle velocity falls in 0.2 m/s intervals covering the range from 2.0 to 5.0 m/s. The direction of particle velocity is mainly upward in the core region, and mainly downward in the annular region, with a wide spread of both upward and downward instantaneous values in the transition zone between the core and annular regions. The downward flow fraction, i.e. the fraction of the measured particle velocities which are downward, is also plotted in this figure. The positive velocities in the core region indicate upward flow of solids. With increasing r/R, the time-mean particle velocity gradually decreases and finally becomes negative. An overall flow map for the reactor section is shown in Fig. 9. At the top of the stripper section, i.e. port X8, it
Fig. 9. Overall distribution of time-average vertical velocity component in reactor section. Base conditions, U f =0.74 m/s, U s=0.25 m/s, G sf=18.6 kg/m2 s. Feed jet penetration calculated from Eq. (2).
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133
Fig. 10. Boundary between upward and downward velocity in reactor section compared with penetration of feed nozzles and jets into the column.
is unclear whether or not there is a wall zone with net downward velocity, although there is definitely a somewhat higher upward velocity in the central region. The wall region at this level is subject to flow reversal, with solids traveling intermittently upward and downward, resulting in extensive turbulence. The gas leaving the stripper section is well dispersed throughout the crosssection by the baffles, with few large bubbles or voids. As shown in Fig. 10, the ratio of the area of the downwardparticle-velocity region to the total cross-sectional area increases with height in the reactor section, reaching a maximum about half-way up the column and decreasing towards the top, probably due to turbulence caused by higher superficial gas velocity there. The average dimensionless radial position (r/R) of this boundary is 0.77, suggesting that the downward-particle-velocity region occupies about 40% of the cross-sectional area on
Fig. 12. Effect of external solids circulation flux on particle velocity distribution measured at the port between feed rings 3 and 4 with U g=0.48 m/s at this level.
average. The radial positions of the nozzle tips and the jet penetration, calculated from Merry [9], are also shown in Fig. 10. The feed jets are seen to penetrate through the transition zone between the core and annular regions. Particles descending in the annular region are entrained by the jets into the upward flowing core region. While ascending, some particles transfer from the core to the wall region where they mainly descend along the column wall. The effect of nozzle insertion distance was investigated by retracting all feed nozzles from an insertion distance of 29 mm to be flush with the wall. Results from measurement port Y4 (between feed rings 3 and 4) are shown in Fig. 11. It can be seen that the annular region where particles descend became narrower, while the radial profile of time-mean particle velocity became less non-uniform over the cross-section. The effect of the solids circulation rate on the time-mean particle velocity profile was investigated by varying G s from 0 to 30.3 kg/m2 s at a level between feed rings 3 and 4. Results are shown in Fig. 12. The cross-sectional average
Table 2 Superficial velocities in reactor section
Fig. 11. Effect of nozzle insertion distance for G s=24.4 kg/m2 s at measurement port Y4 between feed rings 3 and 4 with U g=0.48 m/s at this level.
Commercial baseline
Lower feed rates
Higher feed rates
U f at top of reactor
0.74 m/s
Measurement port Y2 between feed rings 1 and 2 Measurement port Y4 between feed rings 3 and 4
0.63 m/s
0.52 m/s (30%) 0.45 m/s (29%)
0.98 m/s (+32%) 0.81 m/s (+29%)
0.38 m/s (21%)
0.59 m/s (+23%)
0.48 m/s
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Fig. 13. Impact of feed rates through nozzles on radial voidage profiles in reactor. For superficial gas velocities see Table 2.
downward particle velocity caused by the external circulation through the riser is 0.05 m/s for the base conditions where G s=24.4 kg/m2 s at this level. Increasing the solids circulation flux is seen to have caused a slight widening of the annular region of the bed in which the net movement is downward. 3.4. Impact of lower and higher feed rate The impact of the superficial gas velocity on the solid flow was investigated by varying the gas flow through each of the feed jets by ~F40% compared with the base conditions. The superficial gas velocities at the measurement ports are compared in Table 2. Note that the magnitudes of the total changes are less than F40% because
the flow from the stripper remained unchanged. Radial profiles of time-mean voidage determined from ports Y2 (between feed rings 1 and 2) and Y4 (between feed rings 3 and 4) are shown in Fig. 13. At both levels and for all radial positions, the bed became more dilute at higher nozzle feed rates and denser at lower feed rates. The shapes of the profile are similar, with a dense annular region surrounding a more dilute core. The radial profiles of standard deviation of local voidage in Fig. 13 indicate similar variation across the cross-section. Time-mean particle velocity profiles are shown in Fig. 14. The boundary separating upward and downward particle mean motion did not change significantly with the total gas flow rate. However, the particle velocity in the core region did depend on the gas flow, especially in the upper region.
Fig. 14. Impact of feed rates through nozzles on solids flow in reactor. For superficial gas velocities see Table 2.
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The jet penetrations calculated from Merry’s equation are also indicated in Fig. 14. In the upper part of the column, as shown in Fig. 14(a), the particle velocity in the central region decreased dramatically and the radial distribution became flatter when the system was operated at lower gas flow rates, presumably because of reduced jet penetration. The internal circulation between the core and annular regions decreased with decreasing superficial gas velocity.
Uf Ug Us VP Z Z*
4. Conclusions Axial and radial distributions of voidage in the reactor section of a geometrically and dynamically scaled fluid coker cold model were measured by an optical fibre probe. The results show a dense annular region surrounding a more dilute core region. The voidage profiles indicate that the voidage increases smoothly from the wall to the axis of the column, rather than showing a sharp transition between annular and core regions. By measuring the particle velocity in the reactor section with another optical fibre probe, it was confirmed that particles tend to rise in the central core, while descending in the outer annular region. The relative radial position where the time-mean vertical velocity component is zero changes with height. For the tapered column (geometrically similar to industrial reactors) used in this work, the downward-flow annular region occupied ~40% of the overall cross-sectional area. Particles descending in the annular region were entrained by the feed jet into the upward-flowing core region. The boundary between upward and downward velocity regions did not change appreciably when the jet gas flow rates were varied by ~F40% from the base conditions. However, the time-mean particle velocity and internal circulation between the core and annular region did change significantly over this range.
Nomenclature P3 q ðq q Þgdp Ar Archimedes number, ¼ g P l2g D diameter at top of reactor section, m d inner diameter of nozzle, m 0 P dp mean particle diameter, m F cumulative mass fraction under U2 Fr Froude number, ¼ g dPf p g acceleration due to gravity, 9.81 m/s2 Gs local net solids circulation flux corrected for crosssectional area, kg/m2 s G sf net solids circulation flux based on cross-sectional area at top of reactor, kg/m2 s H expended bed depth in reactor section, m L horizontal jet penetration, m t time, s T total sampling duration, s u0 feed nozzle jet velocity, m/s
Greek e l qo qg qp
135
superficial gas velocity at dense phase bed surface, m/s local superficial gas velocity corrected for gas flowrate and cross-sectional area, m/s superficial gas velocity at top of stripper section, m/s vertical particle velocity, m/s height coordinate measured from top of highest stripper shed, m dimensionless height coordinate (height coordinate divided by total height of reactor section) letters local voidage gas viscosity, Pa s gas density at nozzle, kg/m3 gas density in fluidized bed, kg/m3 particle density, kg/m3
Acknowledgement The authors are grateful to Syncrude Canada for sponsoring this work and for permission to publish the results. We also wish to thank D. Famulak, J. Xu and W. Wei for their assistance with experiments and Chevron Canada for providing the FCC particles. References [1] B. Knapper, F. Berruti, J.R. Grace, H.T. Bi, C.J. Lim, Hydrodynamic characterization of fluid bed cokers, in: J.R. Grace, J. Zhu, H. de Lasa (Eds.), Circulating Fluidized Bed Technology VII, C.S.Ch.E., Ottawa, 2002, pp. 263 – 270. [2] J.M. Matsen, Scale-up of fluidized bed process: principle and practice, Powder Technol. 88 (1996) 237 – 244. [3] M.R. Gray, T. Le, W.C. McCaffrey, F. Berruti, S. Soundararajan, E. Chan, I. Huq, Coupling of mass transfer and reaction in coking of thin films of Athabasca vacuum residue, Ind. Eng. Chem. Res. 40 (2001) 3317 – 3324. [4] X.T. Bi, H.P. Cui, J.R. Grace, A. Kern, C.J. Lim, D. Rusnell, X.Q. Song, C. McKnight, Flooding of gas–solids counter-current flow in fluidized beds, Ind. Eng. Chem. Res. 43 (2004) 5611 – 5619. [5] X.T. Bi, J.R. Grace, C.J. Lim, C. McKnight, Hydrodynamics of the stripper section of fluid cokers, 2004, in press. [6] B.A. Knapper, Experimental studies on the hydrodynamics of fluid bed cokers, MSc thesis, University of Saskatchewan, Saskatoon, Canada, 2000. [7] L.R. Glicksman, M.R. Hyre, P.A. Farrell, Dynamic similarity in fluidization, Int. J. Multiph. Flow 20S (1994) 331 – 386. [8] P.A. Farrell, M.R. Hyre, L.R. Glicksman, Importance of the solid-togas density ratio for scaling fluidized bed hydrodynamics, in: L.S. Fan, T.M. Knowlton (Eds.), Fluidization IX, Engineering Foundation, New York, 1998, pp. 85 – 92. [9] J.M.D. Merry, Penetration of a horizontal gas jet into a fluidized bed, Trans. Inst. Chem. Eng. 49 (1971) 189 – 195. [10] J. Copan, Macroscopic modeling of a fluid bed coker and experimental studies of one- and two-phase feed jets, MSc thesis, University of Saskatchewan, Saskatoon, Canada, 1999. [11] A. Donald, H.T. Bi, J.R. Grace, C.J. Lim, Penetration of single and multiple horizontal jets into fluidized beds, Fluidization XI, Ischia (2004) 171 – 178.
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[12] J. Liu, J.R. Grace, X.T. Bi, Novel multifunctional optical-fiber probe: I. Development and validation, AIChE J. 49 (2003) 1405 – 1420. [13] J. Liu, J.R. Grace, X.T. Bi, Novel multifunctional optical-fiber probe: II. High-density CFB measurements, AIChE J. 49 (2003) 1421 – 1432. [14] H.T. Bi, N. Ellis, I.A. Abba, J.R. Grace, A state-of-the-art review of gas–solid turbulent fluidization, Chem. Eng. Sci. 55 (2000) 4789 – 4825. [15] N. Ellis, C.J. Lim, J.R. Grace, H.T. Bi, K.S. Lim, Frequency analysis of pressure fluctuations in turbulent fluidized beds, in: J.R. Grace,
J. Zhu, H. de Lasa (Eds.), Circulating Fluidized Bed Technology VII, C.S.Ch.E., Ottawa, 2002, pp. 287 – 294. [16] H.T. Bi, J. Zhou, S.Z. Qin, J.R. Grace, Annular wall layer thickness in circulating fluidized bed risers, Can. J. Chem. Eng. 74 (1996) 811 – 814. [17] J.R. Grace, Contacting modes and behaviour classification of gas– solid and other two-phase suspensions, Can. J. Chem. Eng. 64 (1986) 353 – 363.