Ind. Eng. Chem. Res. 2001, 40, 5415-5420
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Liquid Vaporization in a Fluidized Bed Karine Lecle` re,† Ce´ dric Briens,*,† Thierry Gauthier,‡ Je´ roˆ me Bayle,‡ Maurice Bergougnou,† and Pierre Guigon§ Department of Chemical and Biochemical Engineering, University of Western Ontario, London, Ontario N6A 5B9, Canada, Institut Franc¸ ais du Pe´ trole, BP3 69390 Vernaison, France, and Universite´ de Technologie de Compie` gne, BP529 60205 Compie` gne Cedex, France
Feedstock injection in the riser is a critical step of the fluid catalytic cracking process. However, although a few models have been developed, no experimental study of the vaporization of droplets in a hot fluidized bed has been published. Experimental studies were thus conducted. Agglomeration of wet particles greatly increases the evaporation time. The first step was thus the study of the different agglomeration processes which may occur when liquid droplets are injected into a hot fluidized bed. There are two kinds of agglomeration. Discrete, round agglomerates are formed when droplets are larger than a critical value, which depends on the bed temperature. Large, flat agglomerates are formed when the amount of liquid injected per unit area is larger than a critical value, which is a linear function of the bed temperature. This implies that the height of the bed remains the same when the bed temperature varies. Experiments were then conducted under agglomerate-free conditions in the case of nonpenetrating droplets. In this limiting case, droplets remain within a few millimeters of the bed surface and contact only a restricted number of particles. An estimation of the vaporization time was obtained by deconvolution of the acquisition signal. Introduction Fluid catalytic cracking (FCC) has been used for decades in order to convert heavy petroleum feedstock into gasoline and other valuable products. Feedstock is introduced at the bottom of the riser, contacts the hot catalyst particles, vaporizes, and is converted to smaller molecules in a few seconds. Fast vaporization results in higher yields. Several models have been developed to describe heat transfer and vaporization in risers. Two types of heat transfer can be considered. In homogeneous vaporization, heat is transferred by convection from the surrounding gas. In heterogeneous vaporization, heat is also transferred by direct conduction between droplets and hot catalyst particles. Homogeneous vaporization has been the object of numerous studies because it concerns a wide range of applications. Buchanan1 developed a model for homogeneous vaporization in the case of the FCC process. The first step is preheating of the liquid to its boiling temperature. The heating time is obtained from the following heat-transfer equations:
theating ) -τheating ln
(
Tg - T Tg - Ti
)
τheating ) FlCpldl/6h The second step is the actual vaporization, at the boiling temperature. Buchanan1 estimated the time required for complete vaporization with an empirical correlation. A correction factor was also introduced to account for the reduction in heat transfer caused by the * Corresponding author. E-mail:
[email protected]. † University of Western Ontario. ‡ Institut Franc ¸ ais du Pe´trole. § Universite ´ de Technologie de Compie`gne.
evolving vapor. The total vaporization time was defined as the sum of these two values (heating and vaporization). The time required for heating the liquid was found to be negligible compared to the vaporization time. Martin2 developed a model using both “d2 law” and “infinite conductivity” models for homogeneous vaporization of droplets. Mass- and heat-transfer equations were written for the vaporizing droplet. By applying the model to a typical industrial case, Martin2 showed the strong effect of the droplet diameter on the time required for complete vaporization. For heterogeneous vaporization, Martin2 considered two different cases depending on whether there was direct contact between droplets and catalyst particles. For droplets large enough to incorporate particles, heat transfer was expected to occur by conductivity. The vaporization process is assumed to proceed until vaporization is complete or until the maximum number of particles has been incorporated inside the droplet. Buchanan1 also developed models for heterogeneous vaporization. He used two different approaches that represent two limiting cases for the vaporization time. First, he considered vaporization by direct contact between particles and liquid droplets. The droplet was assumed to be a solid sphere contacting particles through elastic shocks. The heat-transfer rate was supposed to be infinitely high: the droplet reached thermal equilibrium with every particle it encountered. The time for heating and vaporizing a droplet thus depended on the frequency with which it collided with particles. The second model from Buchanan1 considered the case of heat transfer through a gaseous film, assuming that the Leidenfrost effect occurred. The model equations that were used for homogeneous vaporization were corrected to account for the presence of solid particles. Mirgain et al.3,4 proposed some new ideas to describe the interactions between particles and droplets in the case of heterogeneous heat transfer. Models developed
10.1021/ie001149m CCC: $20.00 © 2001 American Chemical Society Published on Web 10/20/2001
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Ind. Eng. Chem. Res., Vol. 40, No. 23, 2001
by Mirgain et al.4 showed that the catalyst stream had to be of the right voidage. Voidages which were either too low or too high prevented fast, complete vaporization. Voidages of 70-95% were required for fast and complete vaporization. Mirgain et al.4 considered three simplified types of collisions between droplets and particles. As a particle covered with liquid impacted a fresh particle, it could either (a) stick to the new particle, leading to a small agglomerate including both particles and liquid (referred to as impact with adhesion), (b)share the liquid with the new particle, leading to two particles moving separately and both covered with half of the liquid initially involved, or (c) transfer the totality of the liquid to the new particle. In each case, the amount of liquid vaporized between each impact was obtained through a classical heat balance. The droplet diameter was found to be a dominant factor in determining whether complete vaporization could be achieved. This critical diameter depended on the type of collision considered. The highest value for the critical size was obtained with the model of impact with adhesion. As shown by these models, the interaction between droplets and particles is a crucial point for a correct understanding of the vaporization process in a fluidized bed. The droplet diameter appeared to be one of the most important parameters. However, none of the recommendations established from these models had been experimentally confirmed. Although those models represent a good start to understanding the process of vaporization, they might also miss some aspects related to the interaction between droplets and particles that could affect the predictions. For instance, agglomeration can occur when a liquid phase is introduced in a fluidized environment. It is essential to avoid agglomeration as much as possible to keep a reasonable heat-transfer rate between the liquid and the solids and ensure a good vaporization. The process of liquid bridge formation as droplets are injected in a fluidized bed has been studied, and models have been developed to describe fluidized-bed agglomerators. Studies of agglomerators identify conditions that promote agglomeration: they provide information that can be helpful to model unexpected agglomeration. As mentioned in the literature review by Nienow,5 the liquid feed rate needs to remain low enough to avoid problems of wet quenching. Wet quenching is very similar to the problems of agglomeration that are presented here: for large quantities of liquid, the bed comes to the point of saturation where part of the bed forms a cake. However, no model has yet been developed to predict the maximum feed rate before the local saturation of the bed occurs. Moreover, the relative size of the droplets compared to the solid particles is also important (Nienow5), and large droplets are often related to unexpected agglomerates. Although models exist in the literature that propose equations to estimate the vaporization rate, no experimental validation has been done yet to support them. In addition to the validation of the model, experimental studies also provide more information concerning the limits for safe operating regions. This study aims at understanding the interactions that take place between the droplets and particles in a fluidized bed. The objective is not to exactly mimic the operating conditions of an industrial riser but to try to isolate the various
Figure 1. Experimental setup diagram.
processes that contribute to vaporization. In a dense nonbubbling bed, agitation and turbulence are minimized and the probability of agglomerate formation is maximized. This represents the worst case for agglomeration. Thus, agglomerate-free operating conditions in a nonbubbling bed will likely be agglomeratefree in bubbling, turbulent, or circulating beds. This paper presents the study of two limiting cases: the agglomeration of wet particles, which occurs when either the liquid flux or the droplets are too large, and the contact of droplets with a bed in the absence of penetration. Experimental Setup The experiments were conducted in a nonbubbling fluidized bed (cf. Figure 1). The superficial velocity of the fluidization air was kept at 7 mm/s, which is between the minimum fluidization and bubbling velocities. The fluidized bed was thus kept in a dense, nonbubbling state with a bed voidage of ) 0.45. The bed cross section was rectangular (0.45 × 0.40 m), and its height was around 0.40 m. Electrical heaters placed on the walls maintained the particles at the set temperature. The bed solid was a FCC catalyst with a Sauter mean diameter of 65 µm and an apparent particle density of 1500 kg/m3. The fluidization air was preheated with an online heater upstream of the bed. The liquid used was water. It was contained in a pressurized vessel. The liquid was injected in the bed by means of pressure-type oil burner nozzles. The residence time of the fluidized bed inside the spraying region was kept in the same range as that in industrial risers. To do so, the relative velocity of the two media (fluidized bed and spray) was reproduced with the same order of magnitude, although the moving phase was not the same. In risers, the circulating solid passes through the fixed spraying zone, whereas in the experimental setup, the nozzle slides above the stationary fluidized bed. The time required for the nozzle to travel between two points was measured precisely by two photoelectric diodes. The nozzle velocity was calculated from the ratio of the distance between the two detectors to the measured time lag. The nozzle velocity was around 2 m/s for all experiments. The height between the nozzle aperture and the bed surface was kept constant at 15 cm in the experiments presented here. The pressure applied on the liquid (between 40 and 100 psig) controlled the liquid flow rate. The set of nozzles covered a wide range of nominal capacity. The liquid flow rate could be varied from 0.07 to 5.5 L/mn. The nozzles produced a solid cone spray with an angle of 45°. The droplet size distribution was measured by phase Doppler anemometry as well as the droplet velocity. Because nozzles were operated above the bed
Ind. Eng. Chem. Res., Vol. 40, No. 23, 2001 5417
Figure 2. Measurement of the droplet size distribution.
during the reported experiments, droplet size measurements were also conducted outside the fluidization column as shown in Figure 2 for each flow rate. To get a full description of the spray, the size distribution was determined at various radial positions and vertical distances from the nozzle tip. The Sauter mean diameter was around 100-120 µm for the operating conditions, with a velocity between 5 and 10 m/s at the bed surface level. Vaporization above the bed is negligible because the air in this region is below 30 °C. Consequently, the droplet size measurements indicate the actual distribution of the droplets hitting the bed surface during the vaporization experiments. The agglomerates formed by the larger droplets of the spray were recovered with a mesh placed inside the bed under the injection area. The agglomerates in the central measuring section of known surface area were then weighed. Agglomeration: Critical Size This section studies the effect of liquid droplet size on agglomerate formation. The liquid flux was purposely kept relatively low. Presentation of the Experiments. Preliminary studies6 have shown that large droplets (whose diameter is relatively large compared to the particle diameter) are likely to form agglomerates when introduced into the fluidized bed. Agglomerate formation is affected by the solid temperature. The present study seeks to establish whether there is a critical droplet size below which agglomeration becomes unlikely. The term “agglomerate” as used in this paper refers to liquid-solid agglomerates with liquid bridges that are strong enough to keep the particles together. In this case, the liquid inside the agglomerate cannot contact new hot particles and cannot vaporize quickly. Several assumptions were made. First, it was assumed that the volume fraction of liquid in the agglomerate was equal to the gas volume fraction inside the fluidized bed: the liquid displaced the interstitial gas when injected into the bed. This was verified by weighing the agglomerate just after it was formed and after drying. The second assumption was that only the larger droplets could form agglomerates. If a droplet diameter was larger than the critical diameter, it formed an agglomerate. For each experiment, the total amount of liquid injected per unit area was obtained from the liquid flow rate and the nozzle velocity over the bed. This total amount was compared to the amount of liquid involved in the agglomeration per unit area, as determined by weighing the recovered agglomerates. The volume fraction of liquid forming agglomerates was then used to
Figure 3. Critical diameter as a function of the bed temperature. Table 1. Nozzles A and B Characteristics nominal capacity at 100 psig, gph
water flow rate at the operating pressure, mL/s
30 60
21 42
nozzle A nozzle B
Table 2. Droplet Velocity and Critical Diameter for Three Injection Options injection single drop droplet velocity at the bed surface (m/s) critical droplet diameter (µm)
nozzle B
nozzle A
0.5
4-5
7-8
