Experimental validation of a new heat transfer intensification method for FCC external catalyst coolers

Chemical Engineering and Processing 75 (2014) 19–30

Contents lists available at ScienceDirect

Chemical Engineering and Processing: Process Intensification journal homepage: www.elsevier.com/locate/cep

Experimental validation of a new heat transfer intensification method for FCC external catalyst coolers Xiuying Yao a , Fuwei Sun b , Yongmin Zhang a,∗ , Chunxi Lu a a b

State Key Laboratory of Heavy Oil Processing, China University of Petroleum, Beijing 102249, China China Kunlun Contracting & Engineering Corporation, Beijing 100037, China

a r t i c l e

i n f o

Article history: Received 31 July 2013 Received in revised form 8 October 2013 Accepted 17 October 2013 Available online 26 October 2013 Keywords: Intensification Heat transfer Fluidized bed External catalyst cooler Hydrodynamics FCC

a b s t r a c t In petroleum refining industry, external catalyst cooler is a key device in FCC units processing heavy residue feedstock. In this study, a new heat transfer intensification method was proposed for FCC external catalyst coolers, which aims to increase their bed-to-wall heat transfer coefficient by enhancing the internal solids mixing and thus the particle renewal on heat tube surface by a double-distributor design. To validate this idea, a large cold model with similar heat tube design and heat transfer mechanism to industrial catalyst coolers was built. Heat transfer coefficient and axial profiles of particle concentrations were measured under different operating conditions. The experimental results proved the feasibility of this heat transfer intensification idea. Higher bed-to-wall heat transfer coefficient, smaller fluidizing gas usage and higher adjustment flexibility are realizable in the new catalyst cooler. It is also learned from this study that uniform gas distribution, limited wall effect, good fluidization state are necessary to achieve good heat transfer performance in FCC external catalyst coolers. An effective height was speculated from the axial tube wall temperature distributions, within which the heat transfer intensification of the new catalyst cooler is effective. This effective height is also found to rise with increasing superficial gas velocity. © 2013 Elsevier B.V. All rights reserved.

1. Background In a modern petroleum refinery, a catalyst cooler is an indispensable device in a fluid catalytic cracking (FCC) unit processing heavy resid feedstock (i.e. RFCC unit) [1,2]. It removes the superfluous heat in regenerator exceeding the requirement for unit heat balance by contacting high-temperature catalyst particles with heat transfer tubes with flowing liquid water in a fluidized bed to produce valuable high-pressure steams. There are catalyst coolers installed inside or outside an FCC regenerator, which are named as internal and external catalyst coolers, respectively [3,4]. Internal catalyst coolers were developed early in the 1960s, which were actually multiple coiled or vertical heat exchange tubes immersed in the dense bed of a regenerator. Due to its unchangeable heat transfer area, the adjustment range of an internal catalyst cooler’s cooling capacity is very small, which limits the operating flexibility of an FCC unit. Moreover, once tube leakage happens, the whole FCC unit must be shut down, resulting in great economic loss. Due to these reasons, various kinds of external catalyst coolers were developed and commercialized since 1980s when heavier feedstock and frequently changing feedstock properties became common in many

∗ Corresponding author. Tel.: +86 10 89731269; fax: +86 10 89739017. E-mail address: [email protected] (Y. Zhang). 0255-2701/$ – see front matter © 2013 Elsevier B.V. All rights reserved. http://dx.doi.org/10.1016/j.cep.2013.10.005

petroleum refineries [5,6]. An FCC external catalyst cooler is actually a fluidized bed heat exchanger placed outside a regenerator. Usually, multiple vertical heat tubes are installed. Its cooling capacity can be easily changed by adjusting the solids circulation rate, static bed level or superficial gas velocity. In some external catalyst coolers, the cooling capacity can be easily changed from 0% to 100%. Compared with an internal catalyst cooler, the operating temperature and gas velocity of an external catalyst cooler are usually lower, thus resulting in reduced erosion and higher reliability. According to the flow directions of particles, external catalyst coolers can be categorized into three types, i.e. down-flow, up-flow and back-mixing [4,6,7]. In a down-flow external catalyst cooler, the superficial gas velocity is usually low, i.e. in bubbling or turbulent flow regime, and particles enters from the cooler’s top and leaves from its bottom. In an up-flow external catalyst cooler, the superficial gas velocity is usually higher, i.e. in fast flow regime, and particles enter from the cooler’s bottom, rise by the entraining gas flow and leave from its top. In this type of cooler, heat transfer area cannot change and tube erosion by the fast-moving particles is severe, resulting in both low operating flexibility and reliability. A back-mixing external catalyst cooler is also operated at low superficial gas velocities, but it has only one opening with the connected regenerator. Particles enter and leave the cooler only by the solids mixing mechanism, i.e. the internal solids circulation in a fluidized bed. Its heat transfer area cannot be changed and solids circulation

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water

water

Nomenclature Aw Cw Cpp hw hd hl hr H0 kpa m q Tb Tb Tin Tout Tw Tw1–5 Tw u0 u1

u2

area of heat transfer surface [m2 ] specific heat of water [J/kg] specific heat of particle [J/kg] bed-to-wall heat transfer coefficient [w/(m2 K)] convection heat transfer coefficient of dense phase [w/(m2 K)] convection heat transfer coefficient of lean phase [w/(m2 K)] radiation heat transfer coefficient [w/(m2 K)] static bed height [m] thermal conductivity of particle packet [w/(m K)] water mass flow rate [kg/s] cooling capacity [w] bed temperature [◦ C] average bed temperature [◦ C] inlet temperature of water [◦ C] outlet temperature of water [◦ C] wall temperature of heat tube [◦ C] wall temperature at different height of the heat tube [◦ C] average wall temperature of heat tube [◦ C] superficial gas velocity [m/s] gas volume flow rate from the central perforatedplate gas distributor divided by the bed crosssectional area [m/s] gas volume flow rate from the side pipe gas distributor divided by the bed cross-sectional area [m/s]

Greek symbols ıd time fraction of dense phase contacting with heat tube surface [–] εpa voidage of particle packet [–] p particle density [kg/m3 ]  pa root mean residence time of packets at the heat transfer surface [s]

rate to and from the cooler and regenerator can only be changed in a narrow range, depending mainly on the operating gas velocity. Therefore, its operating flexibility is very low. Due to its better operating flexibility and high reliability, downflow type external catalyst coolers are usually preferred choices for most RFCC unit designs. Fig. 1 shows the schematic of a typical down-flow external catalyst cooler. Except for the shell and gas distributor, effective and reliable heat tube design is the key in external catalyst cooler designs. Currently, most external catalyst coolers choose a concentrically double-tube design as their heat exchange tubes [2,8–10]. As shown in Fig. 1, saturated liquid water first enters in the small center tube and then turns to rise through the annular flow area. Heat transfer begins from the outside hot catalyst particles to the inside liquid water, generating valuable high-pressure steam. The catalyst cooler design shown in Fig. 1 also represents a typical design in various FCC external catalyst coolers employed in China, where each heat tube has independent inlets and outlets [4,8]. As leakage in one tube occurs, this tube can be easily shut down and separated. The cooler can still work without stopping, thus avoiding an unscheduled unit shutdown. Therefore, its reliability can be further improved. In some heat tube designs, vertical fins, e.g. in Fig. 2, are welded onto the heat tube surface to increase the heat transfer area and the cooling capacity [8,11]. Without considering the inside heat transfer, an FCC external catalyst cooler is essentially a fluidized bed with multiple vertical tube internals. Its heat transfer properties, e.g. the cooling

water+steam

water+steam

hot particles air

air

cold particles Fig. 1. Schematic of a typical down-flow external catalyst cooler.

capacity, reliability, and operating flexibility, are closely related to the properties of bed hydrodynamics and solids mixing. Despites of usage for decades in commercial FCC units, problems such as low heat transfer capacity, unstable catalyst circulation and tube damages were frequently reported in industrial catalyst coolers [12–15]. These incidents will reduce the unit throughput due to cooling capacity bottleneck. Sometimes, the entire unit has to be shut down, resulting in serious economic loss. These facts indicate a serious lacking of knowledge in understanding the inside complex flow behavior and its relationship with heat transfer properties, which is the key to achieve optimized design of external catalyst coolers.

Fig. 2. Typical finned heat tube bundle in an external FCC catalyst cooler.

X. Yao et al. / Chemical Engineering and Processing 75 (2014) 19–30

In the last several decades, there were a lot of efforts on improving the equipment reliability and intensifying heat transfer of various external catalyst coolers. Improvement of equipment reliability is mainly related to the structural and manufacturing technologies [14,15]. Heat transfer intensification is to increase a heat tube’s the cooling capacity. For a FCC external catalyst cooler, heat transfer intensification measures involve three categories according to the following universal heat transfer equation, q = hw Aw (Tb − Tw ),

(1)

i.e. (a) increasing heat transfer area, (b) increasing heat transfer coefficient, and (c) increasing differential temperature between heat transfer surface and the fluidized bed. Here, hw and Aw are the total bed-to-wall heat transfer coefficient and heat transfer area, Tb and Tw are the bed temperature and heat tube wall temperature, respectively. Adding vertical fins [8,11] like in Fig. 2 primarily belongs to the first category. An alternative method was proposed by Zhang and Zhang [9] who welded a lot of short nail-shaped steel bars on the heat tube surfaces. This method proved its better heat transfer intensification effect in a commercial test. The heat transfer coefficient of the new heat tube was found to be 60–80% higher than a smooth tube. In an industrial FCC external catalyst cooler, solids circulation rate is often used to adjust its cooling capacity. Increasing solids circulation rate can increase the average bed temperature, the differential temperature of heat transfer and thus the cooling capacity. This method belongs to the third category. For the second category, if hw is to be increased, the bed hydrodynamics and solids mixing properties must be changed, which involves selections of superficial gas velocity, structural design and layout of the heat tubes, gas distributor design etc. This is to create suitable hydrodynamic conditions for better solids-wall contact and bed-to-wall heat transfer. Usually, an FCC external catalyst cooler is high aspect ratio fluidized bed with multiple heat tubes. Previous studies on deep fluidized beds of FCC particles [16–19] have shown that local defluidization and severe gas-bypassing are very possible to happen in these beds. As more tubes are inserted or fins are added, the hydraulic diameter of the bed is decreased, which could create possible conditions for slugging, a typical state of bad fluidization quality. Moreover, the existence of vertical tubes and their fins limits solids mixing [20,21] and thus deteriorates solids-wall contact. These influences are all disadvantageous to bed-to-wall heat transfer. Moreover, tube erosion and damage are also the possible results. Systematic study on the hydrodynamics and the related heat transfer properties is very important to achieve optimized design and operation of a FCC external catalyst cooler. However, to our knowledge, there is still no open literature on this issue. To a certain extent, it is due to propriety protection. However, many external catalyst cooler designs are indeed based on empirical data and existing industrial experiences, far from the optimized design. Recently, we established a large cold model of a down-flow external catalyst cooler, where the heat transfer properties and the inside hydrodynamics were experimentally studied [22]. This external catalyst cooler has similar finned heat tubes as in Fig. 2, designed based on the drawing of an industrial down-flow external catalyst cooler. We also used a similar heat transfer mechanism to measure its bed-to-wall heat transfer coefficients. The hydrodynamics was also measured to analyze its relationship with heat transfer properties [22]. This provided a first attempt for systematic study on the complex hydrodynamics and heat transfer in FCC external catalyst coolers. In this study, a new method to increase the bed-to-wall heat transfer coefficient of the down-flow FCC external catalyst cooler was proposed. To validate its feasibility and advantages, two new

21

catalyst cooler models were built. Their heat transfer properties and hydrodynamics, following the same methods, were measured in the large cold model and compared with those of the previous base cooler model. The obtained understandings are expected to provide helps to improve the design of existing industrial FCC catalyst coolers.

2. Brief review of bed-to-wall heat transfer mechanisms in gas–solids fluidized beds and related studies A FCC external catalyst cooler is a typical gas–solids fluidized bed heat exchanger. Before detailed description of this study, a brief review of bed-to-wall heat transfer mechanisms in gas–solids fluidized beds and related studies on heat transfer intensification are first presented. There are three mechanisms of heat transfer between a gas–solids fluidized bed and its immersed surfaces, i.e. particle convection, gas convection and radiation. In most cases, these three terms are thought to be addable, so the overall heat transfer coefficient in a bubbling fluidized bed is commonly written as [23]:

hw = ıd hd + (1 − ıd )hl + hr

(2)

where, ıd is the time fraction of contact by the dense/emulsion phase, ıd hd is the particle convection component during dense/emulsion phase contact, (1 − ıd )hl is the gas convection component during lean/bubble phase contact and hr is the radiation component. For fluidized beds operating at temperatures less than 500 ◦ C, radiation component is generally less important [24]. Previous measurement results [25–32] suggest that hw is several times higher than that for single-phase gas convection. Superficial gas velocity has a profound influence on hw . As superficial gas velocity increases, hw usually first increases and then decreases slightly after peaking at a critical gas velocity which some researchers found near the onset velocity of turbulent flow regime [32,33]. Particle size also influences hw . Smaller hw s were found in fluidized beds of increasing particle sizes. Except for these factors, temperature, pressure, static bed height, and bed internals are all reported to have influences on hw . To predict the heat transfer coefficient in a fluidized bed, many studies chose empirical correlations [30,31,34,35]. Generally, these methods treat heat transfer in fluidized beds as an analogy to heat transfer in a single-phase fluid. The enhancement of heat transfer in fluidized beds is attributed to the scouring action of solid particles on the heat-resistive gas film, decreasing the effective film thickness. However, Chen [36] compared several different correlations against a set of experimental data for heat transfer in a fluidized bed and found that deviations were in an order of 100%. Therefore, these correlations must be used with considerable cautions in scale-up purposes. In a bubbling fluidized-bed heat exchanger, an immersed heat tube surface is actually covered by dense emulsion phase in most time, only being periodically disturbed by the passage of discrete gas bubbles. This has been proved by experimental findings [25,37]. Therefore, the more mechanistic packet renewal model for fluidized bed heat transfer originally suggested by Mickley and Fairbanks in 1955 [38] has gotten the widest recognitions in the fluidization community. In their model, the emulsion phase covering heat transfer surface and separated by discreet bubbles was defined as particle packets. Bed-to-wall heat transfer in a fluidized bed is dominated by transient conduction between particle packets and the surface during periods when packets cover the heat transfer surface. If a particle packet is assumed to be a pseudo homogeneous

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medium of solids volume fraction (1 − εpa ), Mickley and Fairbanks [38] derived the average heat transfer coefficient of dense phase,

 hd = 2

kpa p cpp (1 − εpa ) pa

1/2 ,

To bag filter gas outlet tube

cyclone

(3)

Here, kpa is packet thermal conductivity and εpa is void fraction in the packets.  pa is the root mean residence time of packets at the heat transfer surface, which is an indicator of the packet renewal at heat transfer surface and also the most important factor influencing heat transfer coefficient of the dense phase. Combining Eqs. (2) and (3), it can be seen that, if radiation is negligible, the two dominant influencing factors are the packet/emulsion fraction ıd and its renewal frequency 1/ pa . However, the two factors are difficult to be determined by experiment and the first layer of particles next to the surface also has different voidage from the bulk emulsion phase. Therefore, Chandran and Chen [39] established a method which can make these model parameters and thus the heat transfer coefficient determined quantitatively. Comparison between the predicted results by their modified model and experimental data in two fluidized beds of different particle sizes showed great agreement. To date, most fluidized bed heat transfer studies dealt with large Geldart B/D particles. For fluidized beds of fine Geldart A particles like in an FCC external catalyst cooler, related heat transfer studies are rare. Stefanova et al. [32] employed an electrically heated vertical copper pipe to measure the bed-to-wall heat transfer coefficient of a fluidized bed of FCC particles. They found that the maximum heat transfer coefficient in the range of superficial gas velocities where the transition to turbulent regime occurred. Moreover, the radial profile of heat transfer coefficient becomes flat after the bed steps into the turbulent flow regime. Using same measurement method, Stefanova et al. [33] investigated the influence of bed scale on bed-to-wall heat transfer in two columns of ID 0.29 m and 1.56 m. The magnitude of the maximum heat transfer coefficients was found to be unaffected by the column diameter, but occurred at higher superficial velocities in the larger column. In turbulent flow regime, the local heat transfer coefficients were independent on radial position in both large and small columns. Bai [40] surveyed and analyzed more than 3000 sets of operating data of commercial FCC external catalyst coolers. The dependence of heat transfer coefficients on main operating conditions was analyzed, finding the most important factors are superficial gas velocity and particle concentration. The fin efficiency was found to be inconstant, varying dependent on operating conditions. A modified heat transfer coefficient correlation was obtained with higher accuracy. In fluidized beds of silica sand of mean particle diameter 307 ␮m and 200 ␮m, Rasouli et al. [41] examined the effectiveness of annular fins on heat transfer coefficient of a horizontal immersed tube. Heat transfer coefficient of circular horizontal finned tube was found to be lower than that of the unfinned horizontal tube but total heat transfer increases because of the larger surface area. 3. Experimental setup and measurement methods 3.1. Experimental setup A schematic diagram of the experimental unit is shown in Fig. 3. A cylindrical fluidized bed of 0.476 m ID and 3 m height was used to model a down-flow FCC external catalyst cooler, where finned vertical heat tubes were installed. In order to keep hydraulic similarity to industrial FCC external catalyst coolers, its hydraulic diameter was kept similar to an existing industrial down-flow external catalyst cooler. Pressured air supplied by a Roots blower was the fluidizing gas. Fluidized particles were FCC equilibrium catalyst of

dipleg Cooled water holding plate heat tube plexiglas fluidized bed column rotameter Hot water gas distributor

surge tank blower Air

Fig. 3. Schematic diagram of the experimental unit.

a mean diameter 69.4 ␮m and particle density 1500 kg/m3 , typical Geldart A particles. In order to steady the gas flow, a surge tank was installed in the pipeline. Air flow rate was controlled by a valve and measured by a rotameter. When the bed was fluidized, smaller particles entrained to the top gas outlet tube were collected by a highefficiency cyclone separator. Collected solids returned to the bed via a dipleg to keep a constant solids inventory in the bed. Smaller uncollected fine particles were captured by a bag filter connected to the cyclone exit. These particles were also dumped into the bed periodically to keep the bed solids material same particle size distribution. In studying the external catalyst cooler of traditional design [22], nine vertical steel finned heat tubes of 76 mm OD and 1.2 m height were distributed in the column. Detailed designs of the tube bundle and the fin structure can be seen in Fig. 4. These tubes had similar external geometry to industrial designs. There were two finned sections of 500 mm height at the upper and downer part of each heat tube. Ten fins of 10 mm width and 2 mm thickness at each finned section were welded around the tube. There was a 100 mm separation distance between the two fin sections, as seen in Fig. 4(b). For the nine tubes, one was fixed in the center of bed and the other eight around a concentric circle of 334 mm in diameter, as seen in its lateral arrangement shown in Fig. 5. A pipe distributor (57 mm OD) with 36 holes of diameter 10 mm, corresponding to an aperture ratio of 1.5%, was installed in the column bottom, 100 mm below the bottom of heat tubes, to distribute the fluidizing gas. This cooler was used as the comparison base for the other two new coolers, so it was called the base catalyst cooler (BCC) in this study. In all experimental runs, the static bed height was maintained to be 1.45 m to guarantee all heat tubes buried in the dense bed. Superficial gas velocity ranged from 0.05 m/s to 0.65 m/s, including both a bubbling and turbulent flow regimes. 3.2. Design of the two new catalyst coolers In view of the convective heat transfer mechanism, bed-to-wall heat transfer in a bubbling fluidized bed is more closely related to the surface solids renewal if the bed gas velocity is kept in a small adjusting range, because its apparent bed density and the induced emulsion fraction will not change too much. In order to intensify

X. Yao et al. / Chemical Engineering and Processing 75 (2014) 19–30

23

draft tube

perforated plate

gas flow solids flow

central tube

Fig. 6. Schematic of a recirculating fluidized bed (redrawn from Yang [42]). Fig. 4. Pictures of the catalyst cooler and finned heat tubes.

solids renewal on the heat transfer surface, strengthening solids mixing should be an effective way. Previous studies [42–45] have shown that a recirculating fluidized bed, i.e. a fluidized bed with two gas inlets and a draft tube (see Fig. 6), can boost macro solids internal circulation and stronger axial solids mixing. Its hydrodynamic mechanism lies in the double gas entrances in the bed bottom. More gas is injected through the bed center, resulting in higher voidage in the bed center and lower voidage near the bed wall. Due to density difference, solids form a circulation pattern shown in Fig. 6, i.e. central up flow and down flow in the annular region near the bed wall. The draft tube, if positioned at a proper height above the central tube, can further increase the solids circulation rate [42]. Based on this idea, we proposed a new method to intensify heat transfer in a FCC external catalyst cooler as shown in Fig. 7. Multiple vertical heat transfer tubes are placed in the top section of the cooler, similar to traditional designs. To guarantee free solids flow through the new external catalyst cooler, two gas distributors are

placed in the bottom. A perforated plate distributor is in the bed center and the majority of gas is injected through it. An annular pipe distributor is placed near the wall, slightly above the central perforated plate distributor. A minority of gas is injected through it. Likewise, density difference makes stronger solids circulation in a similar pattern as shown in Fig. 6. Stronger solids circulation scours the heat tube surface more frequently, enhancing solids renewal. Based on this heat transfer intensification idea, two new catalyst cooler models were manufactured. They have same gas distributor design, but the above heat tube structures were slightly different. The former pipe distributor in BCC was removed and replaced by a double-distributor design as shown in Fig. 7. The diameter of the central perforated plate distributor was 320 mm, placed 200 mm below the bottom of the heat tubes. There were 32 holes of diameter 6 mm, amounting to 0.9% of the total bed column flow area. The

water

water water+steam

water+steam hot particles

air

air air

cold particles Fig. 5. Heat tube arrangement in the BCC.

Fig. 7. Idea of the new FCC external catalyst cooler.

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X. Yao et al. / Chemical Engineering and Processing 75 (2014) 19–30

Tout

heat tube constant-temperature water trough temperature controller

partition plates (a) ACC-1

T Tin

heater

(b) ACC-2

Fig. 8. Heat tube arrangements in the two types of catalyst cooler.

flowmeter

outside pipe (32 mm OD) gas distributor, placed 100 mm above the central perforated plate distributor, had 12 holes of diameter 6 mm, amounting to 0.2% of the total bed column flow area. The function of the pipe distributor was mainly to aerate the down-flowing solids. There were 8 finned heat tubes in each of the two new catalyst coolers. They were called the annular catalyst cooler (ACC) in this study. To avoid interfere on solids flow, the central heat tube was removed as seen in Fig. 8(a) and (b). Later, in order to further promote particles circulation, two sections of vertical partition plates were installed along the heat tube center circle to form a draft tube, as shown in Fig. 8(b). The heights of the two sections of partition plate circles and their separation distance were the same as the two fin sections. In this study, the two catalyst cooler shown in Fig. 8(a) and (b) were named as ACC-1 and ACC-2, respectively. The region among the heat tubes was called as the draft tube region and the region between the heat tube and bed wall was called as the annular region. The aim of these two new coolers was to provide non-uniform gas distribution and promote stronger inner solids circulation. At a constant gas flow rate, their bed-to-wall heat transfer coefficients were expected to higher than that of the BCC due to the stronger solids mixing and surface solid renewal. Moreover, the adjustable range of their cooling capacity can be further increased by adjusting the gas flow rate ratio from the two gas distributors.

external pump Fig. 9. Schematic diagram for determination of the heat transfer coefficient.

cool water

Tw5 Tw4 Tw3 Tw2 Tw1 hot water

3.3. Measurement methods

Fig. 10. Relative positions of the five tube wall thermocouples with respect to the two fin sections.

In this study, the most important measurement parameter is the bed-to-wall heat transfer coefficient. In order to guarantee measurement accuracy, a heat transfer mechanism analogy to industrial catalyst coolers as shown in Fig. 9 was employed. Hot water from the constant-temperature trough was pumped into a heat tube of the catalyst cooler by an external pump. Trough water temperature was kept constant, usually in the range of 70–90 ◦ C, by a temperature controller. The trough had a large volume of 100 L to make temperature control easier. Hot water was pumped into the heat tube from its bottom inlet and then returned into the trough to recover its lost heat. During its passing through the heat tube, heat transferred from the hot water, through the steel tube wall and the fins, to the cold fluidized particles in the bed. Two Pt-100 pole thermocouples were installed at the inlet and outlet of the heat tube to measure the inlet and outlet water temperatures (Tin and Tout ), respectively. The bed-to-wall heat transfer coefficient (hw ) can be calculated by the following heat balance equation, Cw m(Tin − Tout ) = hw Aw (Tw − Tb )

(4)

Here, Cw , m, Aw are the specific heat of water, the calibrated water mass flow rate and the heat transfer area, respectively. The heat

transfer area, Aw , only accounts for the tube outer surface area, without including the added fin surface area. Average tube wall temperature (TW ) was calculated by the temperatures measured by five Pt-100 chip thermocouples adhered at same pitches on the surface of the heat tube. Table 1 lists their axial positions and Fig. 10 shows their relative positions with respect to the two fin sections. The average bed temperature (Tb ) was also determined by five Pt-100 pole thermocouples, which were inserted at same axial positions but different radial positions of the fluidized bed, to make data more representative. According to previous studies [46,47], the measured heat transfer coefficient mainly includes the contributions of gas and solids convections. This is similar to the heat transfer mechanism in an industrial external catalyst cooler where its operating temperature is usually below 500 ◦ C and heat transfer by radiation is less important [24,40]. A small difference is the heat transfer direction. In industrial catalyst coolers, heat transfers from hot particles to saturated water in the heat tube. However, in this study, heat transfers from the inside hot water to the cold fluidized particles. Before each heat transfer coefficient measurement, we waited about 20–30 min to get the bed temperature to be stabilized. Each

X. Yao et al. / Chemical Engineering and Processing 75 (2014) 19–30

25

Table 1 Axial positions of the five tube wall thermocouples. Num.

1

2

3

4

5

Distance above the bottom pipe distributor (mm)

140

440

740

1040

1340

measurement period lasted for 3 min. During measurement, the bed temperature was almost constant. The maximum deviation between the five bed temperatures was less than 2 ◦ C in most tests. In view of heat balance, some will be lost from the small tubes connecting the inlet and outlet of the measured heat tubes. Moreover, some of these surface areas were not accounted in the measured heat transfer coefficient. These two factors were the main measurement error sources. However, due to their small diameters (20 mm OD) and short distances, these surface areas were much smaller than the transfer area of the measured heat tube. Moreover, these errors existed in all the three measured catalyst coolers. Therefore, the final conclusion on the comparison among the three catalyst coolers will not change. Fig. 11 shows a set of typical heat transfer coefficient measured in BCC. It can be seen that heat transfer coefficient first increases and then decreases with increasing gas velocity, peaking at a critical superficial gas velocity u0 = 0.4 m/s. If the cross section area occupied by the heat tubes is subtracted, the actual transitional gas velocity is 0.49 m/s that approximates to the onset turbulent fluidization velocity computed by the correlation of Cai et al. [48]. The value range of hw is generally agreeable with previous studies of Stefanova et al. [32,33] in similar fluidized beds and the industrial data [9,40], demonstrating the reliability of the measurement in this study. According to the packet renewal theory, this can be explained as the following: for u0 < 0.4 m/s, the frequency of particle renewal on tube surface plays a dominant role, which increases with increasing gas velocity; for u0 > 0.4 m/s, the increasing frequency of particle renewal cannot counterbalance the effect of decreased bed density, so the heat transfer coefficient decreases with increasing superficial gas velocity. In our previous study [22], the effect of inlet water temperature on the measured heat transfer coefficient was studied. Three inlet water temperatures, 60 ◦ C, 75 ◦ C and 90 ◦ C were tested. It was found that slightly higher hw s were measured under higher inlet water temperatures, but the increments were not remarkable, usually within 10%. This may be related to the enhanced gas convection under high temperatures. In order to make data comparable, all the heat transfer coefficients for the three catalyst coolers in this

study were measured at a same inlet temperature, 90 ◦ C. Moreover, higher heat transfer coefficient was observed in the center heat tube than in the heat tubes near the bed wall. We attributed it to the stronger wall effect which suppressed the particle renewal on their tube surfaces. Except for the heat transfer coefficient, axial profiles of particle concentrations were also measured in this study, with which we tried to correlate the different heat transfer properties in different catalyst coolers. The axial profiles of particle concentration were determined from the bed pressure gradient measured by pressure transducers. We installed 9 gauge pressure transducers along the bed height, which could determine 8 apparent particle concentrations. The gauge pressure transducers were provided by Sailing Technology, Beijing, China, which had a range of 0–35 kPa and a measurement accuracy of 0.2% of its full scale. Each transducer was connected to a short tube of 4 mm ID. Steel fiber was inserted into its tip that connects the bed. For every test, a measurement period of 60 s under a sampling frequency of 200 Hz was used to obtain more reliable data. 4. Results and discussion 4.1. Heat transfer coefficient Fig. 12 shows the measured bed-to-wall heat transfer coefficients in ACC-1. Note here, u2 in Fig. 12 is defined as the gas volume flow rate of the pipe distributor divided by the whole column cross-sectional area. Likely, u1 is defined as the gas volume flow rate of the perforated plate distributor divided by the whole column cross-sectional area. The superficial gas velocity of ACC-1 or ACC-2 is the sum of u1 and u2 , i.e. u0 = u1 + u2

(5)

To facilitate comparison, the curve of hw vs. u0 of BCC is also drawn in Fig. 12, i.e. the red dashed curve. It can be seen that a similar trend of heat transfer coefficient as that in BCC is observed when u2 is small, i.e. u2 < 0.031 m/s. The peaking superficial gas velocity is very close to that of BCC. However, a double peak trend appears in the hw vs. u0 curves when u2 exceeds 0.039 m/s. The highest heat

550 500

550

450

500

hw, w/(m2.K)

2

hw, w/(m .K)

600

400 o

Tin=90 C

350

H0=1.45m

450 400 350

BCC

300

ACC-1 u2=0.016 m/s u2=0.023 m/s

250

u2=0.031 m/s

u2=0.039 m/s

u2=0.047 m/s

u2=0.055 m/s

300 250

0.1

0.2

0.3

0.4

0.5

u0, m/s Fig. 11. Typical heat transfer coefficient data measured in BCC.

0.6

200

0.1

0.2

0.3

0.4

u0, m/s

0.5

0.6

Fig. 12. Measured bed-to-wall heat transfer coefficients for ACC-1.

0.7

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X. Yao et al. / Chemical Engineering and Processing 75 (2014) 19–30

600

600

550

550 u1=0.078 m/s

450 400 350 300

u1=0.164 m/s

500

hw, w/(m2.K)

hw, w/(m2.K)

500

450 400 350

1

2

3

4

300

5

2

4

6

u1/u2 (a) 600

600

550

550 u1=0.336 m/s

450 400

12

450 400 350

350 300

10

u1=0.508 m/s

500

hw, w/(m2.K)

hw, w/(m2.K)

500

8

u1/u2 (b)

5

10

15

20

25

300

10

15

20

u1/u2

25

30

35

u1/u2 (d)

(c) Fig. 13. Effect of u1 /u2 on heat transfer coefficient.

transfer coefficients appear near u0 = 0.25 m/s, far smaller than that in BCC. At a given superficial gas velocity, hw increases continuously with increasing u2 . However, this effect is more pronounced when u2 is small. The heat transfer intensification effect of ACC-1 can be readily observed except at two small u2 s, i.e. u2 = 0.016 m/s and u2 = 0.023 m/s. Very likely, bad fluidization states exist near the column wall when the gas flow rate from the pipe distributor is too small, resulting in less effective solids aeration. Even worse, local defluidization may exist. The expected strong solids internal circulation as shown in Fig. 7 may not happen, which suppresses both local and macro solids mixing in the bed, affecting negatively the particle renewal on the tube surface and bed-to-wall heat transfer. As more gas is introduced into the bottom wall region, there appear two effects that may improve heat transfer. One is that gas distribution become evener, which is in favor of good fluidization quality. The other is the possible induced stronger solids circulation as shown in Fig. 7 can further enhance solids renewal on tube surface. Especially, the heat transfer intensification effect of ACC-1 is more pronounced at smaller superficial gas velocities. This means smaller fluidizing gas usage in ACC-1 if similar cooling capacity is demanded. The remarkable effect of u2 on hw also provides ACC-1 more flexibility in adjusting its cooling capacity. At a certain superficial gas velocity, the change of the gas flow rate ratio from the perforated plate and pipe distributors, i.e. u1 /u2 , can change hw and the cooling capacity flexibly in a large range. Fig. 13 shows the effect of u1 /u2 on hw at several typical u1 s. As u1 /u2 decreases, hw s all show considerable increases, which provide a 30–50% adjusting range. For some FCC external catalyst coolers whose bed level control is not available or solids circulation rate control is not flexible, this is undoubtedly a simple and effective revamp option.

Table 2 Measured heat tube surface temperatures at different heights and the corresponding inlet and outlet water temperatures (u1 = 0.078 m/s). u2 (m/s)

Tin

Tw1

Tw2

Tw3

Tw4

Tw5

Tout

0.016 0.023 0.0031 0.039 0.047 0.055

88.82 88.88 89.17 89.22 89.52 89.97

74.84 75.21 73.59 73.77 73.32 73.92

71.59 71.62 67.94 68.07 65.60 65.91

72.56 72.36 68.99 68.75 64.75 65.46

72.83 72.78 71.07 70.52 69.11 70.31

70.34 70.30 70.15 70.09 70.21 70.92

80.81 80.85 81.53 81.60 82.34 82.90

Table 3 Measured heat tube surface temperatures at different heights and the corresponding inlet and outlet water temperatures (u1 = 0.164 m/s). u2 (m/s)

Tin

Tw1

Tw2

Tw3

Tw4

Tw5

Tout

0.016 0.023 0.0031 0.039 0.047 0.055

90.18 90.12 90.02 90.04 90.15 90.20

79.20 77.33 76.43 76.70 76.37 76.79

76.87 74.48 73.70 72.56 72.11 72.68

73.69 71.52 69.37 68.52 68.38 69.18

75.14 74.03 73.11 72.91 73.36 73.25

72.28 71.73 71.44 71.66 71.92 72.18

83.23 83.05 82.79 82.82 82.90 83.07

During the measurement, we found an interesting phenomenon, as can be explained by Tables 2–5. For all heat transfer coefficient measurements in ACC-1, the tube wall temperature always decreased with increasing height, i.e. in a sequence as Tw1 > Tw2 > Tw3 > Tw4 > Tw5 , although sometimes higher temperature gradients were found in the bed bottom zone [22]. However, at a certain small superficial gas velocity, there is a lowest tube wall temperature appeared in the middle of the tube wall. For example, the lowest temperature at u1 = 0.078 m/s appears between Tw2 and Tw3 , i.e. 440–740 mm height above the bottom pipe distributor, as

X. Yao et al. / Chemical Engineering and Processing 75 (2014) 19–30 Table 4 Measured heat tube surface temperatures at different heights and the corresponding inlet and outlet water temperatures (u1 = 0.250 m/s). Tin

Tw1

Tw2

Tw3

Tw4

Tw5

Tout

0.016 0.023 0.0031 0.039 0.047 0.055

88.79 88.25 87.31 86.64 86.17 85.88

78.17 74.90 72.37 72.64 71.83 71.41

73.73 69.43 67.17 67.50 67.28 66.52

70.91 66.07 65.52 65.01 66.25 65.06

69.61 64.89 64.87 64.57 65.67 64.79

66.40 64.95 65.01 65.15 65.59 64.92

80.37 79.82 79.18 78.41 77.68 77.41

Table 5 Measured heat tube surface temperatures at different heights and the corresponding inlet and outlet water temperatures (u1 = 0.336 m/s). u2 (m/s)

Tin

Tw1

Tw2

Tw3

Tw4

Tw5

Tout

0.016 0.023 0.0031 0.039 0.047 0.055

89.00 88.53 88.34 88.17 88.03 87.90

78.93 76.31 74.56 74.24 73.72 74.04

73.49 69.19 68.57 67.71 68.11 67.95

71.07 66.97 66.77 66.47 67.05 66.95

69.66 65.72 65.87 65.84 66.65 66.52

67.07 64.51 64.87 65.03 65.20 66.07

80.72 80.30 80.21 80.11 80.20 80.29

shown by the bold values in Table 2. Moreover, this lowest temperature point seems to rises as u2 and the total superficial gas velocity increase. This trend can be further proved in Table 3 at higher superficial gas velocities. All the lowest wall temperatures are located at Tw3 , a higher position. At u1 = 0.250 m/s, there are still several such lowest temperature points, but they rises to a higher positions at Tw4 (see Table 4). When u1 exceeds 0.336 m/s, no such lowest temperature points can be found. Why such a strange phenomenon appears? We speculate that there exists a specific height range within which the heat transfer intensification of ACC-1 is effective and this specific height increases with increasing superficial gas velocity. Certainly, this speculation should be further improved by finer experimental or numerical studies. Fig. 14 shows the measured bed-to-wall heat transfer coefficients in ACC-2. All the curves of hw vs. u0 show a similar single-peak trend as that in BCC, which is different from ACC-1 at higher u2 s. Their peaking superficial gas velocities are also near that of BCC. Increasing gas flow rate from the side pipe distributor will also enhance heat transfer in ACC-2. Its heat transfer intensification effect can also be observed in some operating conditions. It is most pronounced when u2 is highest, i.e. at u2 = 0.055 m/s. Likely, heat transfer intensification effect is more pronounced in smaller superficial gas velocities. Clearly, its intensification effect is inferior to ACC-1, which was not expected during our experimental

550 500

hw, w/(m2.K)

450 400 350

BCC

300

ACC-2 u2=0.016 m/s u2=0.023 m/s

250

u2=0.031 m/s

u2=0.039 m/s

u2=0.047 m/s

u2=0.055 m/s

200

0.1

0.2

0.3

0.4

u0, m/s

0.5

0.6

Fig. 14. Measured bed-to-wall heat transfer coefficients for ACC-2.

600 560 520

hw, w/(m2.K)

u2 (m/s)

27

480 440 BCC ACC-1 ACC-2 u2=0.055 m/s

400 360 320 280

0.1

0.2

0.3

0.4

0.5

0.6

u0, m/s Fig. 15. The measured heat transfer coefficients for ACC-2.

design stage. We attribute this to the stronger wall effect in ACC-2. As seen in Fig. 8, the space between heat tubes and the bed wall is small. The minimum distance between the tube surface and the bed wall is only about 80 mm. As shown in Fig. 8(b), the existence of fins and addition of partition plate further enhance the wall effect, which may suppress both axial and radial solids mixing and thus the solids renewal on heat tube surface. As stated in our previous study [22], due to stronger wall effect, the heat transfer coefficient of the side tube is lower than that of the central tube. The stronger wall effect may also affect negatively the heat transfer intensification effect of ACC-2. Otherwise, the effective intensification height above mentioned in ACC-1 was also observed in the tube wall temperature data in ACC-2. Similar trend with superficial gas velocity was also observed, further proving our previous speculations. Fig. 15 compares the best heat transfer performance of the three types of catalyst cooler, which enables the intensification effect of the annular catalyst coolers more clearly observed. For the highest hw s, ACC-1 and ACC-2 are 9.5% and 4.7% higher than BCC, respectively. The superior heat transfer intensification effect of ACC-1 to ACC-2 can be more clearly observed in this figure. Otherwise, the stronger heat transfer intensification effect in smaller superficial gas velocities in ACC-1 and ACC-2 also indicates smaller fluidizing gas usage and potential energy saving in industrial units. 4.2. Axial profiles of particle concentration Fig. 16 shows the effect of u1 , i.e. the gas flow rate from the central perforated plate distributor, on the axial profiles of particle concentration in the two annular catalyst coolers shown in Fig. 8. In the two typical operating conditions, u2 is kept constant and at relative low values. In this study, the gas flow from the central perforated plate distributor contributes a dominant fraction of the whole fluidizing gas stream. Therefore, its change can cause remarkable variance in the axial profiles of apparent particle concentration in both ACC-1 and ACC-2. An interface between the dense bed and freeboard can be clearly observed from these profiles. Particle concentration in the dense bed keeps relatively constant, but a sharp decrease is found with increasing bed height. It is found that the apparent particle concentration in the dense bed decreases with increasing u1 , while the apparent particle concentration in the freeboard increases with increasing u1 . The addition of partition plate in ACC-2 does not show significant influence on the effect of u1 on axial profiles of particle concentration. It seems from Fig. 16(b) that the effect of u1 on the particle concentration of the dense bed is stronger in ACC-2. An obvious increase in apparent

750

Apparent particle concentration, kg/m

Apparent partcle concentration, kg/m

3

X. Yao et al. / Chemical Engineering and Processing 75 (2014) 19–30

3

28

600 450 300

u=0.016 m/s 2

150

u=0.078 m/s 1

u=0.164 m/s 1

u=0.250 m/s 1

u=0.336 m/s 1

0

u=0.422 m/s 1

u=0.508 m/s 1

0.0

0.5

1.0

1.5

2.0

750 600 450 300

u2=0.031 m/s

150

u1=0.078 m/s

u1=0.164 m/s

u1=0.250 m/s

u1=0.336 m/s

0

u1=0.422 m/s

u1=0.508 m/s

0.0

2.5

0.5

1.0

1.5

Height, m

Height, m

(a) ACC-1

(b) ACC-2

2.0

2.5

2.0

2.5

3

Apparent particle concentration, kg/m

Apparent particle concentration, kg/m

3

Fig. 16. Effect of u1 on axial profiles of particle concentration.

750 600 u=0.164 m/s 1 450

u=0.016 m/s 2 u=0.023 m/s 2

300

u=0.031 m/s 2 u=0.039 m/s 2

150

u=0.047 m/s 1 u=0.055 m/s 2

0

0.0

0.5

1.0 1.5 Height, m

2.0

750 600

u1=0.164 m/s u2=0.016 m/s

450

u2=0.023 m/s u2=0.031 m/s

300

u2=0.039 m/s 150

u1=0.047 m/s u2=0.055 m/s

0 0.0

2.5

0.5

1.0

1.5

Height, m

(a) ACC-1

(b) ACC-2

Fig. 17. Effect of u2 on axial profiles of particle concentration.

750 600 450 300 150

BCC

u0=0.190 m/s

ACC-1 u0=0.180 m/s ACC-2 u0=0.195 m/s

0 0.0

0.5

1.0

1.5

2.0

2.5

Apparent particle concentration, kg/m

3

Apparent particle concentration, kg/m

Fig. 17 shows the effect of u2 , i.e. the gas flow rate from the side pipe distributor, on the axial profiles of particle concentration in the two annular catalyst coolers. Due to the relative contribution to the total fluidizing gas stream, increasing u2 does not significant influence in both ACC-1 and ACC-2. Fig. 18 compares the axial profiles of particle concentration of the three catalyst coolers in similar superficial gas velocities. As a whole, there is no big difference between BCC and the two annular catalyst coolers. However, the particle concentration in the dense

3

particle concentration is observed near the dense bed surface in both ACC-1 and ACC-2. This effect is caused by the heat tubes, which possessed ca. 20% of the total bed cross sectional area. The abrupt increase of flow area above the top of heat tubes can cause a significant decrease in actual bed gas velocity, which causes an abrupt increase in the local apparent particle concentration. This phenomenon did not happen at u1 = 0.508 m/s in ACC-2, as shown in Fig. 16(b), which may be related to the decreased bed level due to higher particle entrainment.

750 600 450 300 BCC 150

u0=0.400 m/s

ACC-1 u0=0.352 m/s ACC-2 u0=0.367 m/s

0 0.0

0.5

Height, m

(a) Fig. 18. Comparison of axial profiles of particle concentration.

1.0

1.5

Height, m

(b)

2.0

2.5

X. Yao et al. / Chemical Engineering and Processing 75 (2014) 19–30

bed of BCC seems to be slightly lower than those of ACC-1 and ACC2, as seen in both Fig. 18(a) and (b). This is partially related to that there was one more heat tube in BCC, resulting in higher actual gas velocity and lower particle concentration in BCC. However, no direct relationship with their intensified heat transfer can be concluded from the axial profiles of particle concentrations in ACC-1 and ACC-2. Finally, an objective analysis can be drawn from this study that, except for the intensification effect caused by the strengthened solids circulation by the double-distributor design shown in Fig. 7, the improved heat transfer performance in the two new annular catalyst coolers may also be related to their improved gas distribution. After all, the single pipe distributor in BCC cannot realize same gas distribution uniformity and fluidization quality in ACC-1 and ACC-2. It is a little unexpected that the performance of ACC-2 is inferior to ACC-1, which may be related to its stronger wall effect that suppresses particle renewal on heat tube surface. Therefore, a good design of an external catalyst cooler should take a balance among good fluidization quality and solids mixing intensity. It can thus be expected that the current empirical guideline on selecting hydraulic diameter and fin arrangement should be further optimized to achieve better heat transfer performance in industrial catalyst cooler designs. More finely designed fundamental studies are still needed, which is also our following research plans. 5. Conclusions After above-mentioned experimental studies, at least the following conclusions can be drawn: (1) The intensification heat transfer effects of the annular catalyst cooler are partially validated. Higher bed-to-wall heat transfer coefficient, smaller fluidizing gas usage and higher adjustment flexibility are realizable in the new annular catalyst cooler. (2) To achieve good heat transfer performance in FCC catalyst coolers, uniform gas distribution, limited wall effect, good fluidization state are necessary, which are related to the optimization in selecting appropriate hydraulic diameter and fin arrangements. (3) An effective height was speculated from the axial tube temperature distributions, with which the heat transfer intensification of the new catalyst coolers is effective. This effective height rises with increasing superficial gas velocity. (4) More finely designed fundamental studies are needed to further discover the relationship between the hydrodynamics and its intensified heat transfer. Acknowledgements The authors acknowledge the financial supports by the National Natural Science Foundation of China (21276273), the Ministry of Science and Technology of China (2012CB215004 and 2012BAE05B02), the Ministry of Education of China by the Program for New Century Excellent Talents in University (NCET-11-0733) and the Science Foundation of China University of Petroleum, Beijing (KYJJ2012-03-11). References [1] T.Y. Chan, D.S. Soni, F. Zhang, Advances in a catalyst cooler technology, Petrol. Technol. Q. Autumn 87 (1999) 83–84. [2] R. Pillai, P.K. Niccum, FCC catalyst coolers open window to increased propylene, in: Grace Davison FCC Conference, Munich, 2011. [3] Z.P. Lai, Catalyst cooling techniques in heavy oil FCCU, Petrol. Ref. Eng. 25 (6) (1995) 44–48 (in Chinese). [4] J.W. Chen, Catalytic Cracking Process and Engineering (2ED), China Petrochemical Press, Beijing, China, 2005, pp. 1338–1343 (in Chinese).

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