PEPT visualisation of particle motion in a tapered fluidised bed coater

Journal of Food Engineering 93 (2009) 324–336

Contents lists available at ScienceDirect

Journal of Food Engineering journal homepage: www.elsevier.com/locate/jfoodeng

PEPT visualisation of particle motion in a tapered fluidised bed coater F. Depypere a,b,*, J.G. Pieters b, K. Dewettinck a a

Faculty of Bioscience Engineering, Department of Food Safety and Food Quality, Laboratory of Food Technology and Engineering, Ghent University – UGent, Coupure links 653, B-9000 Gent, Belgium Biosystems Engineering, Ghent University – UGent, Coupure links 653, B-9000 Gent, Belgium

b

a r t i c l e

i n f o

Article history: Received 12 June 2008 Received in revised form 15 January 2009 Accepted 31 January 2009 Available online 25 February 2009 Keywords: Radioactive particle tracking Particle circulation Gas–solid fluidisation Conical Drying Coating

a b s t r a c t Detailed knowledge of particle motion is essential in better understanding fluidised bed coating processes. This paper reports on the use of Positron Emission Particle Tracking (PEPT) in order to visualise and quantify the features of the powder movement inside a laboratory-scale GPCG-1 fluidised bed coater (Glatt GmbH, Germany) with a tapered (conical) chamber. In all cases, a clear circulating motion, upwards in the centre and downwards along the walls, was revealed. This circulating pattern lends itself to operations such as particle coating, where controlled application of the coating onto the particles is essential. In contrast to the fluidisation patterns observed in cylindrical fluidised beds or spouted beds without a draft tube, radial movement was stronger and occurred at all heights in the powder bed. By introducing a binary nozzle and increasing the nozzle atomisation air pressure, the circulating motion of the fluidised particles in a tapered vessel was found to be accelerated. When spraying a coating solution downwards onto the fluidised powder particles, effects on the particle motion, depending on coating solution properties and process parameters, were observed. Besides visualisation, PEPT was also used to derive quantitative information such as bed height and particle circulation time, as well as to investigate the size of the coating zone. Generally, it was demonstrated that PEPT can be successfully used to characterise the particle motion in a tapered fluidised bed under both drying and coating conditions. Ó 2009 Elsevier Ltd. All rights reserved.

1. Introduction Film-coating of powders is an added-value technique whereby a spray process is used to encapsulate a pure active ingredient or a mixture of ingredients within a protective material or system (a coating). In general, possible benefits of microencapsulation are controlled release, protection of the core ingredient, increase of the overall product quality and increase of the processing convenience (Depypere et al., 2003). Among the available equipment used for coating or agglomeration in the food, pharmaceutical and chemical industry are fluidised beds, rotating drums and spouted beds (Watano et al., 1994; Dewettinck and Huyghebaert, 1999; Jono et al., 2000; Teunou and Poncelet, 2002; Depypere et al., 2003; Degrève et al., 2006; Werner et al., 2007). With respect to fluidised beds, no other microencapsulation technology can apply as broad a range of coating materials (aqueous solutions, aque-

* Corresponding author. Address: Faculty of Bioscience Engineering, Department of Food Safety and Food Quality, Laboratory of Food Technology and Engineering, Ghent University – UGent, Coupure links 653, B-9000 Gent, Belgium. Tel.: +32 9 264 61 67; fax: +32 9 264 62 18. E-mail address: [email protected] (F. Depypere). 0260-8774/$ - see front matter Ó 2009 Elsevier Ltd. All rights reserved. doi:10.1016/j.jfoodeng.2009.01.042

ous latex dispersions, hot melts – fats, fatty acids, waxes – or organic solvent solutions) (Vilstrup, 2001; Gouin, 2004). Besides coating, the fluidised bed can also be used for drying or agglomeration purposes. In all these unit operations, the fluidisation air serves not only to support the powder bed and maintain the solids circulation, but also acts as the heating or cooling medium. For coating and agglomeration, a spray nozzle is inserted in the fluid bed. Depending on the position of the spray nozzle relative to the powder bed, three batch fluidised bed configurations exist: top-spray, bottom-spray (or ‘‘Wurster” (Wurster, 1959)) and tangential- (or rotor-) spray, with specific advantages and disadvantages for each configuration (Depypere et al., 2003). For reasons of higher versatility (both aqueous and hot melt coatings can be applied), relatively high batch size and relative simplicity of disassembly and cleaning, the top-spray configuration (Fig. 1) is widely regarded as the most appropriate configuration for successful use in the food industry (Dewettinck and Huyghebaert, 1999). As food powders are mostly characterised by a wide particle size distribution, the use of a tapered (conical) fluid bed (cross-sectional area increasing with height) is preferred over a cylindrically shaped vessel. Due to the angled walls of the product container, a more vigorous circulation can be induced in the fluidised bed. A higher gas velocity towards the base keeps larger particles in motion, and a

F. Depypere et al. / Journal of Food Engineering 93 (2009) 324–336

Fig. 1. Scheme of the top-spray fluidised bed coating process and equipment.

lower gas velocity higher up avoids elutriation of the smaller particles (Singh et al., 1992; Jones, 1994). In order to understand the hydrodynamics of any fluidised bed process, detailed knowledge of the particle motion inside the equipment is of crucial importance. However, measuring local hydrodynamics within real processing equipment has remained a major challenge due to a lack of suitable measurement techniques. It is only during the last decades that techniques have been developed to fill this scientific gap. Cheremisinoff (1986), Grace and Baeyens (1986), Yates and Simons (1994), Simons (1995), Werther (1999) and Tayebi et al. (2001) consider some of the options. As the fluidised bed under investigation is characterised by high solids volume fractions, optical visualisation of particle motion inside the equipment is problematic. An earlier paper by Depypere et al. (2005a) described the possibility to determine the expanded bed height using pressure and temperature measurements along the bed wall. Even with transparent walls, (high speed) photographic techniques yield merely information about the behaviour in a thin layer near the wall, which may be far from representative of the behaviour in the interior (Werther, 1999). Techniques such as laser Doppler anemometry (LDA) and particle image velocimetry (PIV) require transparent equipment and bed materials to allow flow behaviour to be quantified. Magnetic resonance imaging (MRI) is largely limited in its scale of operation and requires that experimental equipment cannot contain metal parts: this is a serious limitation for solids motion studies inside real lab-scale and pilot-plant equipment. Non-intrusive measurements can also be performed using tomographic or tracer techniques. With the latter, detectable tracer particles with a hydrodynamic behaviour representative of the suspension material are used to study the particle motion (Tayebi et al., 2001). Tracer experiments have so far been widely used and can successfully provide accurate particle motion data (Baeyens and Geldart, 1986; Grace and Baeyens, 1986). The advantage of a tracer technique over tomography is that Lagrangian type data is provided as the tracer is tracked in 3D space and time (Barigou, 2004). Positron Emission Particle Tracking (PEPT) is a state-of-the-art tracer technique for particle research, being one of the few available non-invasive methods able to visualise and quantify the particle motion. PEPT was invented in the 1980’s at the University of Birmingham as a refinement of Positron Emission Tomography

325

(PET) (Hawkesworth et al., 1986; Bemrose et al., 1988) and its performance is continuously being improved (Parker et al., 1993, 1994, 1997, 2002). PEPT has so far been applied to characterise the particle motion in, for example, mixers, rotating drums, stirred vessels, millers, extruders, heat exchangers, hoppers, vibrated beds, gas–solid and liquid–solid fluidised beds and spouted beds (Forster et al., 2000). For cylindrical beds loaded with particles in Geldart’s ‘‘group B” (Geldart, 1973), Stein et al. (2000) reported PEPT results on the solids motion during bubbling fluidisation. Also, PEPT was used to study the particle motion in the Wurster process operated under drying conditions (Fitzpatrick et al., 2003). More recently, Schaafsma et al. (2006) used PEPT to study the particle flow pattern and granule segregation in tapered fluidised bed granulators. In this work, the PEPT technique developed at the University of Birmingham was used to investigate the particle motion in a laboratory-scale Glatt GPCG-1 (Glatt–Powder–Coater–Granulator) top-spray fluidised bed coater (Glatt GmbH, Germany). The first objective of this study was to obtain a better understanding of the overall solids motion in a tapered fluidised bed. Therefore, this paper reports on the major fluidised bed characteristics of interest: bed height, particle circulation time, velocity profile and residence time distribution. The second objective was to investigate how the particle motion was affected by the presence of the nozzle and the spraying of an aqueous solution of different coating materials downwards onto the fluidised powder bed.

2. Materials and methods 2.1. Fluidised bed experiments The Glatt GPCG-1 (Glatt GmbH, Germany) is one of the most versatile used laboratory-scale fluidised bed coating units, with processing chambers well representative for applications in the food and pharmaceutical field. The device consists of a stainless steel (2.5 mm thickness) tapered vessel (140 mm base diameter, 8.1° taper angle) with a stainless steel woven wire mesh air distributor (172  36 wires/inch) at the base. More information on the equipment dimensions and the distributor used, is given elsewhere (Depypere et al., 2004, 2005a). The fluidisation air flow rate was measured under normal conditions using a rotating vane probe (Testo, Belgium) with a diameter equal to the inlet air duct dimension. Most of the fluidisation experiments were performed for a lower (55 m3/h) and a higher (97 m3/h) level of the fluidisation air flow rate, F, corresponding to a superficial air velocity across the distributor of 1.0 and 1.75 m/s, respectively. At these velocities, the distributor pressure drop was 51 and 110 Pa, respectively. These inlet air conditions were chosen from the air flow rate range representative for normal processing conditions in the current fluidised bed equipment. Hereby, gas–solid fluidisation is performed in the bubbling fluidisation regime, at gas velocities 5–10 times the minimum fluidisation velocities of the powders used (Depypere et al., 2005b). The temperature of the fluidisation air entering the plenum was set at 75 °C. In all experiments, 750 g of bulk material was introduced into the bed. The following bulk materials were tested: glass MicrobeadsÒ (Sovitec, Belgium), sucrose/starch beads (‘‘non-pareils”) (Penwest Pharmaceuticals, U.S.A.) and microcrystalline cellulose (MCC) beads ‘‘Cellphere” (FMC Biopolymer, Belgium). Using a laser diffraction device (Mastersizer S equipped with a MSX-64 Dry Powder Feeder, Malvern, UK), a 300 mm lens (0.5–900 lm) and a 1000 mm lens (4–3500 lm), the particle size distribution of the investigated powders was measured and a surface weighted average diameter, d32, and volume weighted average diameter, d43, were obtained (10 replicates). Particle density, qp, was measured

326

F. Depypere et al. / Journal of Food Engineering 93 (2009) 324–336

via toluene pycnometry at 25 °C (five replicates). Particle shape was analysed using a Leitz Diaplan light microscope equipped with a Nikon Coolpix 4500 digital camera. In Table 1, a detailed overview of the characteristics of all investigated powders is given. Based on the values for d32 and particle density, all investigated powders could be classified as Geldart type B powders and are, consequently, expected to fluidise easily and bubble at gas velocities above the minimum for fluidisation (Geldart, 1973). In most of the experiments and unless stated otherwise, a binary nozzle was centrally mounted at 121 mm above the distributor. This position is further referred to as the ‘low’ position while the nozzle can also be mounted in the unit at 168 mm above the distributor, i.e., the ‘high’ position. In a binary nozzle, coaxial atomisation air is used to shear a central exiting coating liquid stream into droplets. In this way, the coating liquid is atomised and a hollow cone of droplets is formed. Under spraying conditions, the atomisation air pressure of the binary nozzle was varied from 2 to 4 bar. When the nozzle was not active, which is e.g., the case during preheating of particles prior to coating, maintenance atomisation air pressure (1.5 bar) was applied to avoid particles obstructing the nozzle outlet. In the coating experiments, two different coating materials were investigated: sodium caseinate and gelatine hydrolysate, both prepared as an aqueous (10–15 wt% dry matter) solution. For the binary nozzle operating at 3 bar pressure, different coating experiments were set up. In each test, the non-pareils were allowed to preheat for 20 min (i.e., a ‘drying’ experiment in the presence of a nozzle under maintenance conditions) before initiating the spraying of the coating solution. In a first test, a 10 wt% sodium caseinate solution was sprayed onto the fluidised powder bed at a liquid spray rate of 5 g/min. In the following three tests, gelatine hydrolysate was used as the coating material. One test was performed with equal settings as for the one with sodium caseinate (10 wt%, 5 g/min). In another test with gelatine hydrolysate the liquid spray rate was increased (10 wt%, 7 g/min) while in the last test both the liquid spray rate and the solution concentration were increased (15 wt%, 7 g/min). 2.2. PEPT technique Using PEPT, a single tracer particle labelled with a radioisotope (in this case fluorine-18, 18F) was introduced into the system of interest. Upon decay of the radioisotope, positrons were released which annihilated with neighbouring electrons and a pair of back-to-back c-rays was produced. Using a number of successively detected c-ray pairs the tracer was located in three dimensions using triangulation. The PEPT technique is illustrated in Fig. 2 and is described more in detail by Parker et al. (1993, 2002). With the current developments in the PEPT camera, timing electronics, location algorithms and tracer activities, a tracer moving at 1 m/s can typically be detected to within 0.5 mm 250 times per second (Parker et al., 2002). However, reasonable tracking can still be obtained for tracer speeds up to 10 m/s (Barigou, 2004). Recently, multiple-particle tracking was proposed as a further improvement of positron particle tracking (Yang et al., 2006, 2007). For this research, the Glatt GPCG-1 fluid bed device was positioned between the two camera detectors, having a useful cross-

Fig. 2. PEPT principle: c-ray detection and triangulation.

sectional area of 500  400 mm2 and being separated from each other by 609 mm. The region of interest – the product container and the expansion chamber – was situated within the borders of the detection window. 2.3. Tracer production for PEPT Interpretation of PEPT measurements relies on the tracer motion being representative of the bulk or of an intended fraction of the bulk. Also, the experiment duration must be long enough to allow the tracer to visit all possible locations in the experimental device sufficiently many times that good statistics are obtained for the process of interest. This in turn demands the use of a tracer with a sufficiently long half-life that it can still be detected accurately at the end of the experiment. With the Birmingham Scanditronics MC40 Cyclotron, radioactive tracers were produced using the surface adsorption method (Fan et al., 2006a, 2006b). Firstly, activated water was produced through direct bombardment with a 33 MeV 3He beam. This converted some of the 16O in the purified water to fluorine-18, 18F, a radioisotope with a half-life of about 110 min. Subsequently, this freely ionic 18F could be transferred to the selected tracer particle via surface adsorption (Forster et al., 2000). The extent to which the tracer particle is representative for the bulk material is of paramount importance for the PEPT measurements. For glass beads and microcrystalline cellulose, a particle with a mean particle size was selected from the bulk powder and activated through surface adsorption. More realistic food powders such as non-pareils, however, dissolve in water. As a consequence, the surface adsorption technique could not be directly applied to one of these particles. Based on the similarities in hydrodynamic properties (Table 1), microcrystalline cellulose (MCC) was selected as tracer for non-pareils. An MCC particle with a size equal to the

Table 1 Core and tracer particle properties (mean ± standard deviation). Powder

d32 (lm)

d43 (lm)

Density (kg/m3)

Shape

Glass beads Non-pareils MCC *

196.54 ± 0.64 345.00 ± 1.29 425.12 ± 0.84

204.18 ± 0.56 385.45 ± 0.76 452.13 ± 0.82

2467 ± 3 1459 ± 10 1527 ± 4

Spherical Oval, porous Nearly spherical

*

Microcrystalline cellulose.

F. Depypere et al. / Journal of Food Engineering 93 (2009) 324–336

mean size of the bulk material and a shape close to that of the bulk powder particles was selected. Using the surface adsorption technique, a sufficiently high radioactivity could be obtained for all labelled particles. Before each experiment, the initial radioactivity level of the tracer was registered. After applying the PEPT algorithm and averaging the data for all experiments (each lasting 1200 s), it was found that the glass beads were detected to within 0.87 mm 55 times per second. An even better location accuracy and temporal resolution was obtained for the MCC beads that were detected to within 0.53 mm 87 times per second. During top-spray fluidised bed coating experiments, an aqueous solution of coating material was sprayed onto the fluidised powder bed. Since the radioisotope fluorine-18, used in this PEPT study to label the tracer particle, is water soluble, it was anticipated that the activity could ‘‘leak” from the tracer particle into the bed resulting in increased background noise and a reduction in the frequency and accuracy of tracer detection. While PEPT is able to cope with a high degree of background noise (Parker et al., 1993), an important aspect of this study was to assess the effect, if any, of loss of activity from the tracer to the bulk particles. 2.4. PEPT data processing For each experiment, the first 20 min of recorded data were used as inputs for the PEPT processing. The raw dataset obtained from the PEPT measurements typically consists of hundreds of thousands of tracer particle X-, Y- and Z-coordinates as a function of time. The ‘‘Track” software, developed at the University of Birmingham, enables visualisation of the solids motion pattern via a pseudo-real time tracer trajectory, occupancy plots and velocity vector plots. Techniques were also developed to extract further quantitative information from the location data: the expanded bed height, the total circulation time and the frequencies of particles entering a specific zone.

3. Results and discussion 3.1. Particle motion in a tapered vessel The effects of air flow rate and aspect ratio AR (bed height/ diameter) on the solid circulation patterns in cylindrical fluidised beds of Geldart B particles are summarised by Kunii and Levenspiel (1991). At low air flow rates and for beds with AR close to, but less than, unity, upflow along the wall and downflow at the bed axis were reported. At higher air flow rates, however, rising bubbles, which are the primary cause of solids motion in fluidised beds (Geldart, 1986) cause a reversal of this flow pattern to upflow in the centre and downflow along the wall (Lin et al., 1985; Baeyens and Geldart, 1986). For a cylindrical bed of glass beads (AR < 1) and a bed of lactose particles (AR  1), Maronga and Wnukowski (1998) observed both migration patterns for the same fluidising gas velocity. For a tapered bed, Kim et al. (2000) observed a downward flow of finer particles along the walls from radial sampling measurements. For a bench-scale conical fluidised bed dryer, Tanfara et al. (2002) identified a centralised lean core flow and a denser annular region, with flow patterns similar to those in a spouted bed (Mathur and Epstein, 1974). It should be pointed out that the taper angle of the conical fluidised bed used in the work of Tanfara et al. (2002) (±20°) was well above that of the equipment used in this research (8.1°). For the Glatt GPCG-1 fluidised bed operating under drying conditions, Chaplin et al. (2004) reported a more uniform fluidisation over the whole cross-section of the bed. By means of PEPT experiments performed over 20 min, in this work, time-averaged information as a function of location was

327

obtained. For glass beads fluidised at an air flow rate of 97 m3/h, Fig. 3 shows time-averaged PEPT data for a period of 1200 s. The data are presented as combined occupancy and velocity vector plots in the XY- and ZY-centre plane, respectively, averaged over 20 mm thickness in the orthogonal direction. The scale represents the occupancy, defined as the fraction of the total run time during which the tracer was located in each 5  5 mm2 bin of the volume considered. It should be noted that occupancy is not the same as residence time as no distinction is made between few longer visits and many shorter ones (Hoomans et al., 2001; Barigou, 2004). Generally, the bin size for velocity calculations is chosen to be larger than that used for occupancy calculations. The instantaneous particle velocity was obtained using the six-point method (Parker et al., 1995) in which a weighted average value over a series of five sequential locations is calculated. The magnitude and direction of the tracer average velocity in each 10  10 mm2 bin are shown as an arrow originating from that bin, with a length proportional to the velocity magnitude. The average velocity within an individual bin is not shown if less than three values of velocity have been assigned to that bin. Similar, time-accumulated location information can be shown, e.g., in a graph in which all locations visited by the tracer are marked with a point. For the first 1200 s of the experiments, Fig. 4 shows the accumulated particle location plot in the XYand ZY-centre plane, respectively. In these plots, all locations were projected on both centre planes. In this way, the presence of dead zones could be easily determined. Earlier computational fluid dynamics simulation studies (Depypere et al., 2004) revealed that, due to the lateral position of the fluidisation air inlet in the plenum section of the Glatt GPCG-1 device, the air flow pattern above the distributor which is currently used will be asymmetric as the pressure drop over the distributor is low compared to commonly required values (Geldart and Baeyens, 1985). Therefore, higher air velocities at the left side of an XYprojected plane were predicted. PEPT experiments confirmed that even in the presence of particles, asymmetries in the fluidising air above the distributor persisted. In an accumulated particle location plot in the XY-centre plane, an inclined particle bed was clearly noticeable (Fig. 4). From the PEPT studies, it could be confirmed that the fluid bed provided good mixing of the powder mass. The low occupancy region close to the wall on the right hand side of Fig. 4b is due to an observation window projecting into the fluidised bed. The PEPT experimental results in Fig. 3 clearly show an orderly circulating particle motion, upwards in the centre and downwards along the walls. This orderly solids circulation lends itself to operations such as particle coating, where controlled application of the coating onto the particles is essential. While an upwardly directed spray nozzle will reinforce this circulation, it is not clear yet how a downwardly directed spray nozzle will influence the solids motion. In all experiments, however, a clear circulating particle motion was demonstrated, so the fluidisation pattern could not be described as random and unrestricted, as was commonly assumed so far (Jones, 1994). The particle circulation observed in this work is of the same form as that seen in spouted beds and Wurster columns, but with important differences. In spouted beds, which may or may not contain a cylindrical ‘‘draft tube” to guide the flow, a high-velocity, high-voidage ‘‘spout” projects through the bed to the surface. Particles travel up this spout and are then projected into the freeboard, from which they are deposited onto the surface of a downward-moving annulus, which is not itself fluidised. In a Wurster coater, a draft tube is present and again two distinct zones can be distinguished: a central region of high-velocity upward particle motion and a slowly-moving annulus of higher occupancy, which in this case may be close to fluidisation. In the case considered

328

F. Depypere et al. / Journal of Food Engineering 93 (2009) 324–336

Fig. 3. Combined occupancy and velocity vector plots in the XY-centre (left) and ZY-centre (right) plane for glass beads fluidised at F = 97 m3/h.

Fig. 4. Accumulated particle location plot in the XY- (left) and ZY-plane (right) for glass beads fluidised at F = 97 m3/h.

here, the bed is fluidised across its cross-section and upward movement is generated, as in conventional parallel-walled fluidised beds (Stein et al., 2000) by upward-moving bubbles. As a consequence, a gradual transition in occupancy  in both the radial and axial directions – was observed. In an attempt to quantify aspects of the velocity vector plots of Fig. 3, Fig. 5 shows the radial velocity component as a function of radial position for four different axial positions, namely Y = 125, 170, 215 and 260 mm. These axial levels are representative of the bottom-bed, central-bed, top-bed and above-bed regions. Fig. 5 illustrates that comparatively strong radial particle movement occurred at all bed heights. Above the bottom-bed region, there is no evidence that the velocity of radial movement, which occurs from the centre to the wall, changes between the centralbed, top-bed and even freeboard region. As particles move radially towards the walls at all heights above the bottom-bed zone, radial movement is stronger than normally seen in cylindrical bubbling beds (Stein et al., 2000; Fan et al., 2008), in which radial mixing is limited except close to the bed surface and near the gas distributor. The same remark can be made when the radial movement reported in this paper is compared to what is observed for spouted beds without a draft tube (Forster et al., 2000). For the lower air flow rate (F = 55 m3/h), Fig. 6 shows the timeaveraged (total time: 1200 s) combined occupancy and velocity vector plots in the XY-centre plane (DZ = 20 mm) for an experiment with MCC and non-pareils, respectively. It should be noted that in both cases MCC was used as the tracer particle. It was found

that the fluidised bed height of both powders was very similar (120–125 mm), indicative for the similarity in size and density of MCC and non-pareils. On the other hand, the slight difference in shape between the nearly spherical MCC beads and the oval shaped non-pareils became apparent from the comparison of their respective occupancy and velocity vector plot. Whereas for MCC, a smooth circulation extending to the bottom of the bed was observed, non-pareils were found to move very slowly in a zone immediately above the distributor. This region of lower particle exchange is probably attributable to the elongated shape of the non-pareils, hindering the smooth circulation of the particles. Importantly, the observation of this difference in fluidisation pattern between MCC and non-pareil beads when an MCC particle was used as the tracer, in each case supports the use of MCC to follow the motion of non-pareils. This also indicates that the motion of the tracer particle is not only dependent on its hydrodynamic and physical characteristics, but is also to some extent determined by the motion of the bulk particles surrounding the tracer particle. It was therefore concluded that a tracer with a similar size and density as the bulk material can be used with confidence to investigate the particle motion of the bulk material. The same observations were made at the high air flow rate. 3.2. Bed height A procedure was developed to derive the expanded bed height of a fluidised powder mass from the occupancy data of a fluidisation

F. Depypere et al. / Journal of Food Engineering 93 (2009) 324–336

329

Fig. 5. Radial velocity as a function of radial distance and axial position (Y = 0.125 m (j), 0.170 m (h), 0.215 m (N), 0.260 m (4)). Positive values indicate tracer movement from the centre to the walls.

Fig. 6. Combined occupancy and velocity vector plots in the XY-centre plane for microcrystalline cellulose (left) and non-pareils (right) fluidised at F = 55 m3/h.

experiment in the tapered column of the Glatt GPCG-1 device. The basis for the calculation is an XY-centre plane occupancy plot (DZ = 20 mm) for a complete experiment (1200 s), as e.g., shown in the left plot of Fig. 3. Reasons for selecting this narrow volume element are: (a) bed height should be assessed when the particle is moving mainly in the upward direction, which takes place in a central region and (b) outside this central region, the powder is redirected to the walls and may be influenced by wall effects. As could be seen from the 2D plot of Fig. 3, on every height Y, multiple bins (5 mm size) could be distinguished. This occupancy plot was one-dimensionalised by averaging the occupancy values for all bins of the same height. This gives a plot of average occupancy as a function of height in the vessel. For the example discussed (glass beads, higher air flow rate), the 1D-occupancy profile is shown in Fig. 7. Both average and standard deviations are depicted. Also, an indication is given on how the expanded bed height may be derived from a graph of this kind. For all experiments of this study, a similar profile to that in Fig. 7 was observed. In the region immediately above the distributor, located at Y = 107.5 mm, there is a sharp rise in 1D-occupancy with increasing height. However, a maximum value is reached at some level, after which the 1Doccupancy decreases steadily. The following bed height discrimi-

nation criterion was defined: an ‘‘above-bed region” is defined when the average 1D-occupancy falls below an absolute critical value. From reviewing the 2D-occupancy plots of experiments where there was an abrupt change in occupancy at the bed surface level, this critical value was determined as being 0.5. The use of this criterion in the above described procedure proved to be valuable for experiments where the transition in bed density was more gradual, hence where the bed height could not a priori be derived from the original 2D-occupancy plots. For the case shown in Fig. 7, the upper bed level was found at Y = 227.5 mm, so the total bed height was 120 mm. 3.3. Particle circulation time and circulation frequency Both in spouted bed coating and Wurster coating, regular fluidisation patterns are established, so the measurement of circulation time becomes straightforward (Fitzpatrick et al., 2003). In a tapered fluidised bed designed for top-spray coating processes, it was shown above that particles rise to varying heights, making the definition of circulation time difficult. Based on the expanded bed height, the powder bed was divided into three parts: a bottom-section extending from the bottom to

330

F. Depypere et al. / Journal of Food Engineering 93 (2009) 324–336

Fig. 7. Determination of expanded bed height from the 1D-occupancy profile as a function of height in the vessel (glass beads, F = 97 m3/h).

25% of the bed height, a central section between 25% and 75% of the bed height and a top-section above 75% of the bed height, including the freeboard region (Fig. 8). Based on earlier definitions for circulation time in cylindrical fluidised beds (Wong et al., 2001; Stein et al., 2002), the total circulation time was defined as the sum of: (1) the time the tracer spends in the bottom-section, (2) the time during which the tracer moves from the bottom-section to the top-section, (3) the time spent in the top-section and (4) the down-flow time between the top-section and the bottom-section. In order to allow for the finite uncertainty in location, the tracer was only considered to have moved from a section (either entering or leaving) when a specified minimum distance beyond the boundary had been passed. This distance is called ‘‘fuzziness” and was taken as 5 mm, which corresponds to the bin size used in the occupancy plots. For each of the four movements concerned, the mean residence time and the spread on this value were calculated using the residence time distribution function e(t) (Fogler and Brown, 1992). In Fig. 9, the residence time distribution, characteristic of each of the four movements (bottom, up, down, top), is shown for the experiment with glass beads fluidised at a high air flow rate (Fig. 3). In Table 2, some statistical parameters (mean residence

time s, standard deviation rs, number of observations No) of the distributions given in Fig. 9 (total time: 1200 s) are shown. From Table 2, the overall particle circulation time distribution is seen to have a wide spread. This is also illustrated in Fig. 10 where the Y-displacement of the tracer during a randomly selected time interval of 10 s is shown. This is not a regular circulation, hence it was found that a high number of revolutions did not fulfil the adopted definition of a circulation. Application of the definition of circulation time alone therefore would exclude a large fraction of the data. In other words, knowledge of the mean particle circulation time alone does not fully characterise the particle motion during fluidisation in a tapered column. Equally important is the knowledge of how frequent particles moved between the axial borders of the central region of the bed (between 25% and 75% of the bed height), as an indication of the frequency of particle circulations. The mean time between two successive circulations, tc–c, was calculated by dividing the total experimental time by the number of axial movements, which is the average of the number of upward and downward movements. In the case presented in Table 2, tc–c was calculated as 1200 s/ 813 = 1.48 s. Compared to the derived circulation time of s = 0.75 s, about 50% of tc–c was represented by complete circulations (s) and the remainder of tc–c consisted of incomplete circulations. For the whole set of experimental results of this work, tc–c was found to be about twice the value of the total mean particle circulation time, so the conclusions drawn for the example discussed could be generalised. 3.4. Effect of nozzle presence, position and atomisation air pressure

Fig. 8. Sectioning of the fluidised bed in the determination of circulation time s (Hb = bed height).

Fig. 11 shows the time-averaged (total time: 1200 s) flow pattern for non-pareils, fluidised in the absence of a nozzle, in the presence of an ‘‘inactive” nozzle (maintenance conditions) or in the presence of an active nozzle (at atomisation pressures of 2, 3 and 4 bar). From Fig. 11, it can be observed that positioning a nozzle centrally above the fluidised bed and increasing the atomisation air pressure accelerated the circulating particle motion of the non-pareils, which took place in a smaller radial zone between the nozzle atomisation air cone and the vessel wall. Due to an increase in nozzle atomisation air pressure, the motion of the particles in the vicinity of the distributor changed clearly: for an experiment without nozzle or with an ‘‘inactive” nozzle, a zone

F. Depypere et al. / Journal of Food Engineering 93 (2009) 324–336

331

Fig. 9. Residence time distribution of the bottom- (d), up- (s), down- (j) and top-movement (h) during fluidisation of glass beads at an air flow rate of 97 m3/h.

Table 2 Residence time distribution statistics from Fig. 9 (mean residence time s, standard deviation rs, number of observations No). Section

s ± rs (s)

No

Bottom Up Down Top Total

0.19 ± 0.15 0.12 ± 0.15 0.18 ± 0.17 0.26 ± 0.17 0.75 ± 0.32

1954 811 815 902 4482

of low particle exchange was found immediately above the distributor. With an increase in nozzle atomisation pressure, this ‘‘stagnant” zone was eliminated and the circulating motion was extended towards the distributor. Owing to an increase in nozzle atomisation air pressure, the expanded powder bed height was found to increase, as shown in Fig. 12. This is partly the result of the fact that a central part of the conical vessel is occupied by the downward pointing hollow atomisation air cone, which acts as an ‘obstruction zone’ for the incoming fluidisation air. Provided the fluidised bed voidage remains unchanged, vertical expansion of the powder bed is a consequence. Also the amount of atomisation air introduced into the fluid bed contributes to the increase in bed height: the atomisation air flow rate is between 5–7.5 m3/h at atomisation pressures between 2-4 bar (Dewettinck, 1997). In Fig. 12, also the influence of the nozzle presence and atomisation air pressure on the overall mean particle circulation time

and its breakdown in the residence time in the bottom-, up-, down- and top-section is shown. A decrease in overall mean particle circulation time was observed with the insertion and operation of a spray nozzle. This indicates that the circulating particle motion is accelerated, which can be confirmed from the plots in Fig. 11. From a certain atomisation air pressure on (p = 2 bar), s(bottom) shows a tendency to decrease with increasing atomisation air pressure. This can be explained by the elimination of the ‘‘stagnant” zone, whereby particles in the bottom-section are lifted again sooner compared to the situation without a nozzle. A clear effect on s(down) is observed: a 50% decrease, from 0.43 s in the absence of a spray nozzle to 0.22 s with a top-spray nozzle operating at a pressure of 4 bar, was noticed. This reduction in ‘‘falling time” can be addressed to the overall enhanced circulating motion by the insertion and operation of a spray nozzle. Similarly to the observation for s(down), also a decrease in s(up) may be expected. From Fig. 12 and taking the evolution of expanded bed height into account, it becomes clear that indeed, s(up) tends to decrease with the insertion and operation of the spray nozzle, but the effect is less pronounced compared to s(down). This may be explained by the fact that the particles rise in the vicinity of the downwards sprayed atomisation air cone, which might act as a counterforce for rising particles. Nevertheless, this counterforce seems more than compensated for by the overall accelerated circulation. Finally, s(top) seems not affected by the increase in atomisation air pressure. At this point, it is interesting to refer to previous work by Dewettinck and Huyghebaert (1998), who investigated the effect

Fig. 10. Tracer Y-displacement during fluidisation of glass beads at a higher air flow rate (t = 10 s, randomly selected). The horizontal lines represent the 25% and 75% bed height levels. The vertical lines demarcate tracer movements (A–E) which can all be considered to have passed through one complete circulation.

332

F. Depypere et al. / Journal of Food Engineering 93 (2009) 324–336

Fig. 11. Combined occupancy and velocity vector plot in the XY-centre plane of non-pareils fluidised at F = 55 m3/h: influence of the nozzle presence and atomisation air pressure.

Fig. 12. Influence of nozzle presence and atomisation air pressure on the bed height (d, line), mean particle circulation time and its breakdown (see legend).

F. Depypere et al. / Journal of Food Engineering 93 (2009) 324–336

of nozzle atomisation air pressure on the amount of coating mass deposited onto the core particles. It was found that coating efficiency increased at higher nozzle atomisation air pressures, which was mainly attributed to higher initial droplet velocities. However, also an effect on the particle fluidisation pattern was suggested, but could not further be specified. This work confirmed the effect of the nozzle atomisation air pressure on the fluidisation pattern and coating efficiency, in that faster circulations to higher positions in the vessel may contribute to an increase in coating efficiency. In all of above experiments, the nozzle was mounted in the lowest of two possible vertical positions. When the nozzle was mounted in the ‘high’ position and the fluidisation air flow rate was kept constant (i.e., the nozzle is placed 47 mm further away from the powder bed compared to the ‘low’ position of the nozzle), it was confirmed that its presence and operation still influenced the particle circulation. However, the shift in fluidisation behaviour with nozzle insertion and operation was less pronounced for a nozzle placed in the ‘high’ position when compared to the ‘low’ position, e.g., in contrast to the nozzle in the ‘low’ position, the nozzle in the ‘high’ position did not influence the residence time in the bottom-zone. 3.5. Effect of spraying a coating solution and coating type For the coating experiments with sodium caseinate or gelatine hydrolysate as coating material, Fig. 13 shows the combined occupancy and velocity vector plots (XY-centre plane, DZ = 20 mm), obtained through PEPT recording during 20 min in the coating regime. In Fig. 14, the corresponding overall mean and partial particle circulation times are shown.

333

When comparing the results for the coating experiment with sodium caseinate to the experiment with a nozzle that is active (p = 3 bar), but that does not deliver coating liquid droplets (Fig. 12), a good agreement between the respective overall and partial circulation times can be noticed. This indicates that the motion of the non-pareils, established by a combination of the fluidisation and atomisation air flow, was hardly influenced by the sodium caseinate solution and coating. However, applied in the same dry matter content, gelatine hydrolysate is inherently stickier than sodium caseinate. As a consequence, it may be anticipated that the contacts between the fluidised core particles, coated with gelatine hydrolysate, will not be as elastic as between uncoated non-pareils or between non-pareils coated with sodium caseinate. This difference in coating properties probably explains the generally slower overall particle motion for core particles coated with gelatine hydrolysate versus sodium caseinate (Fig. 14). This was particularly noticed for s(bottom) and s(down), which are associated with movements where particles are more close together compared to the upward movement (influence of the drag force) and the unhindered motion in the top-zone of the fluid bed and the freeboard. When the gelatine hydrolysate dry matter content in the solution and/or the spray rate were further increased, zones of higher occupancies between the nozzle tip and the distributor (central bottom-bed) were noticed. Compared to the particle motion of sodium caseinate coated particles where no zone of slow particle motion at the bottom of the fluid bed was observed, this locally retarded particle movement is most likely to be associated with the higher stickiness of gelatine hydrolysate coated particles.

Fig. 13. Combined occupancy and velocity vector plot in the XY-centre plane: influence of coating material (sodium caseinate vs. gelatine hydrolysate), dry matter content (10–15 wt%) and spray rate (5–7 g/min).

334

F. Depypere et al. / Journal of Food Engineering 93 (2009) 324–336

Fig. 14. Mean particle circulation time and its breakdown (see legend): influence of coating material (sodium caseinate vs. gelatine hydrolysate, dry matter content (10– 15 wt%) and spray rate (5–7 g/min).

Fig. 15. Occupancy plot in the XY-plane (DZ = 20 mm), focussing on the coating zone, for glass beads fluidised at F = 97 m3/h, with the nozzle operating at maintenance pressure in contrast to at 3 bar pressure: atomisation air spray cone identification and spray angle determination.

F. Depypere et al. / Journal of Food Engineering 93 (2009) 324–336 Table 3 Influence of the nozzle spray on the particle motion in the fluidised bed: penetration distance and cone angle results as a function of experimental conditions. Powder

Atomisation pressure (bar)

Penetration distance (mm)

Cone angle ac (°)

Non-pareils Non-pareils Non-pareils Non-pareils Glass beads Glass beads

Maint.* (1.5) 2 3 4 Maint.* (1.5) 3

49 61 87 109 54 86

§ 30.5 17.5 12.4 37.1 17.9

§, No clear determination possible. ‘‘Inactive” nozzle operating at maintenance conditions (1.5 bar).

*

3.6. Spray coating zone assessment All of the above shown occupancy plots were constructed using a bin size of 5  5 mm2. By focussing into the region below the spray nozzle tip and by using a bin size of 1 mm2, it was investigated whether the coating zone could be visualised and characterised. As an example, Fig. 15 shows the occupancy plot in the XY-plane (DZ = 20 mm, bin size = 1 mm2) when 750 g of glass beads were fluidised at a high air flow rate. The results of two experiments are shown: the first experiment with an ‘‘inactive” nozzle operating at maintenance pressure (1.5 bar) and the second one with an active nozzle operating at 3 bar pressure. At the right side of Fig. 15, it is illustrated that the hollow atomisation cone could be distinguished from its surroundings. In the plot for the experiment with the nozzle operating at maintenance pressure, a realistic shape for the coverage of the spray pattern was detected (Schlichting and Gersten, 1997). Due to the low atomisation pressure, a theoretical ‘straight’ coverage could not be attained over a longer distance. At 3 bar, however, the atomisation air pressure proved to be high enough to realise a ‘straight coverage’ over the impact distance of the nozzle. From the latter plot, the spray angle could be derived as indicated in Fig. 15. Based on the occupancy data in the zone around the atomisation cone axis, a procedure was developed to also derive the penetration distance of the atomisation air cone, that is the distance over which the injected air affected the particle motion. In Table 3, the results are summarised for fluidisation experiments with non-pareils and glass beads. From Table 3, it can be seen that the influence of the nozzle atomisation spray extended over a considerable distance in the powder bed. Furthermore, the penetration distance was found to increase with increasing atomisation air pressure. While the nozzle influence was relatively low when operating at maintenance pressure, the impact of a nozzle operating at 4 bar pressure was found to almost reach the distributor plate. Interestingly, also the cone angle could be reasonably well quantified. A decrease in cone angle with increasing atomisation air pressure was found. Furthermore, the results were in good agreement with the information provided by the nozzle (Zweistoffdüse Modell 970/S0) manufacturer (Düsen-Schlick, Germany), who mentioned a cone angle ranging from 10°–40°. Based on PEPT data, it was found to be possible to derive the cone angle and atomisation air penetration depth of a spray nozzle during a fluidised bed coating process. As a result, PEPT was found to be suitable to investigate in situ the size of the coating zone. 4. Conclusions Positron Emission Particle Tracking (PEPT) was found to be a powerful technique in the visualisation of particle motion in a tapered laboratory-scale fluidised bed. The penetrative c-rays in-

335

volved in the PEPT measurements enabled to follow with high location accuracy and temporal resolution the motion of a tracer in an opaque system of this kind, even inside actual processing equipment. It was observed that the fluidisation pattern in a tapered fluid bed is far from being random and unrestricted: a circulating pattern of solids movement, upwards in the centre and downwards along the walls, is established. This pattern lends itself to operations such as particle coating, where controlled application of the coating on the particles is essential. This circulating motion is different from that observed in cylindrical beds, spouted beds or Wurster columns, in that the extent of radial motion is much stronger. With the introduction of a centrally positioned downwards spraying binary nozzle and with an increase in nozzle atomisation air pressure, the circulating particle motion in a tapered fluid bed vessel was accelerated. When an aqueous solution of coating material was sprayed onto the powder bed, PEPT results were not compromised by a loss of radioactivity from the tracer to the surrounding particles. Besides a qualitative analysis, PEPT also offers possibilities to extract quantitative information from the camera recordings. This data can either be used as inputs or for validation in the development of numerical or statistical models of the fluid bed process. Overall knowledge of the flow pattern of the core particles contributes to increased process understanding. In the end, this will favour the production of more homogeneous coatings, thus lead to better coating quality. Acknowledgements The authors wish to thank Prof. Jonathan Seville and Dr. David Parker (Positron Imaging Centre, Birmingham University) for providing the PEPT test facilities and Dr. Andy Ingram and Dr. Xianfeng Fan from the same institution for the practical assistance and the interesting discussions regarding this work. Also Benny Lewille (Ghent University) is gratefully acknowledged for his practical assistance to this study. The authors wish to thank the Fund for Scientific Research– Flanders (Belgium) (F.W.O. Vlaanderen) as well as the Scientific Research Committee of the Faculty of Bioscience Engineering of Ghent University for their financial support. The International Division of the Institute of Food Technologists (IFT) is gratefully acknowledged for having awarded this work with the first price in the 2005 George F. Stewart International Research Competition at the occasion of the 2005 IFT Annual Meeting in New Orleans. References Baeyens, J., Geldart, D., 1986. Solids mixing. In: Geldart, D. (Ed.), Gas Fluidization Technology. Wiley, Chichester, pp. 97–121. Barigou, M., 2004. Particle tracking in opaque mixing systems: an overview of the capabilities of PET and PEPT. Chemical Engineering Research and Design 82 (A9), 1258–1267. Bemrose, C.R., Fowles, P., Hawkesworth, M.R., O’Dwyer, M.A., 1988. Application of positron emission tomography to particulate flow measurement in chemical engineering processes. Nuclear Instruments and Methods in Physics Research A 273 (2–3), 874–880. Chaplin, G., Pugsley, T., Winters, C., 2004. Application of chaos analysis to pressure fluctuation data from a fluidized bed dryer containing pharmaceutical granule. Powder Technology 142 (2–3), 110–120. Cheremisinoff, N.P., 1986. Review of experimental methods for studying the hydrodynamics of gas–solid fluidized beds. Industrial and Engineering Chemistry Process Design and Development 25 (2), 329–351. Degrève, J., Baeyens, J., Van de Velden, M., De Laet, S., 2006. Spray-agglomeration of NPK-fertilizer in a rotating drum granulator. Powder Technology 163 (3), 188– 195. Depypere, F., Dewettinck, K., Ronsse, F., Pieters, J.G., 2003. Food powder microencapsulation: principles, problems and opportunities. Applied Biotechnology Food Science and Policy 1 (2), 75–94.

336

F. Depypere et al. / Journal of Food Engineering 93 (2009) 324–336

Depypere, F., Pieters, J.G., Dewettinck, K., 2004. CFD analysis of air distribution in fluidised bed equipment. Powder Technology 145 (3), 176–189. Depypere, F., Pieters, J.G., Dewettinck, K., 2005a. Expanded bed height determination in a tapered fluidised bed reactor. Journal of Food Engineering 67 (3), 353–359. Depypere, F., Pieters, J.G., Dewettinck, K., 2005b. Powder bed behaviour during fluidised bed processing. Minerva Biotecnologica 17 (4), 203–214. Dewettinck, K., 1997. Fluidized bed coating in food technology: process and product quality. PhD dissertation, Ghent University, Gent. Dewettinck, K., Huyghebaert, A., 1998. Top-spray fluidized bed coating: effect of process variables on coating efficiency. Lebensmittel-Wissenschaft und – Technologie – Food Science and Technology 31 (6), 568–575. Dewettinck, K., Huyghebaert, A., 1999. Fluidized bed coating in food technology. Trends in Food Science and Technology 10 (4–5), 163–168. Fan, X., Parker, D.J., Smith, M.D., 2006a. Enhancing 18F uptake in a single particle for positron emission particle tracking through modification of solid surface chemistry. Nuclear Instruments and Methods in Physics Research A 558 (2), 542–546. Fan, X., Parker, D.J., Smith, M.D., 2006b. Labelling a single particle for positron emission particle tracking using direct activation and ion-exchange techniques. Nuclear Instruments and Methods in Physics Research A 562 (1), 345–350. Fan, X., Parker, D.J., Yang, Z., Seville, J.P.K., Baeyens, J., 2008. The effect of bed materials on the solid/bubble motion in a fluidised bed. Chemical Engineering Science 63 (4), 943–950. Fitzpatrick, S., Ding, Y., Seiler, C., Lovegrove, C., Booth, S., Forster, R., Parker, D., Seville, J., 2003. Positron emission particle tracking studies of a Wurster process for coating applications. Pharmaceutical Technology, 70–78 (September). Fogler, H.S., Brown, L.F., 1992. Distributions of residence times for chemical reactors. In: Fogler, H.S. (Ed.), Elements of Chemical Reaction Engineering. Prentice Hall International, New Jersey, pp. 708–750. Forster, R.N., Seville, J.P.K., Parker, D.J., Ding, Y., 2000. Tracking single particles in process equipment or probing processes using positrons. KONA 18, 139–148. Geldart, D., 1973. Types of gas fluidization. Powder Technology 7 (5), 285–292. Geldart, D., 1986. Gas Fluidisation Technology. Wiley, Chichester. Geldart, D., Baeyens, J., 1985. The design of distributors for gas-fluidized beds. Powder Technology 42 (1), 67–78. Gouin, S., 2004. Microencapsulation: industrial appraisal of existing technologies and trends. Trends in Food Science and Technology 15 (7–8), 330–347. Grace, J.R., Baeyens, J., 1986. Instrumentation and experimental techniques. In: Geldart, D. (Ed.), Gas Fluidization Technology. Wiley, Chichester, pp. 415–462. Hawkesworth, M.R., O’Dwyer, M.A., Walker, J., Fowles, P., Heritage, J., Stewart, P.A.E., Witcomb, R.C., Bateman, J.E., Connolly, J.F., Stephenson, R., 1986. A positron camera for industrial application. Nuclear Instruments and Methods in Physics Research A 253 (1), 145–157. Hoomans, B.P.B., Kuipers, J.A.M., Salleh, M.A.M., Stein, M., Seville, J.P.K., 2001. Experimental validation of granular dynamics simulations of gas-fluidised beds with homogeneous in-flow conditions using Positron Emission Particle Tracking. Powder Technology 116 (2–3), 166–177. Jones, D., 1994. Air suspension coating for multiparticulates. Drug Development and Industrial Pharmacy 20 (20), 3175–3206. Jono, K., Ichikawa, H., Miyamoto, M., Fukumori, Y., 2000. A review of particulate design for pharmaceutical powders and their production by spouted bed coating. Powder Technology 113 (3), 269–277. Kim, H.G., Lee, I.O., Chung, U.C., Kim, Y.H., 2000. Fluidization characteristics of iron ore fines of wide size distribution in a cold tapered gas–solid fluidized bed. ISIJ International 40 (1), 16–22. Kunii, D., Levenspiel, O., 1991. Fluidization Engineering. Butterworth–Heinemann, Boston. Lin, J.S., Chen, M.M., Chao, B.T., 1985. A novel radioactive particle tracking facility for measurement of solids motion in gas fluidised beds. AIChE Journal 31 (3), 465– 473. Maronga, S.J., Wnukowski, P., 1998. The use of humidity and temperature profiles in optimizing the size of fluidized bed in a coating process. Chemical Engineering and Processing 37 (5), 423–432. Mathur, K.B., Epstein, N., 1974. Spouted Beds. Academic Press, New York.

Parker, D.J., Broadbent, C.J., Fowles, P., Hawkesworth, M.R., McNeil, P.A., 1993. Positron emission particle tracking – a technique for studying flow within engineering equipment. Nuclear Instruments and Methods in Physics Research A 326 (3), 592–607. Parker, D.J., Hawkesworth, M.R., Broadbent, C.J., Fowles, P., Fryer, T.D., McNeil, P.A., 1994. Industrial positron-based imaging: principles and applications. Nuclear Instruments and Methods in Physics Research A 348 (2–3), 583–592. Parker, D.J., Hawkesworth, M.R., Beynon, T.D., 1995. Process applications of emission tomography. Chemical Engineering Journal 56 (3), 109–117. Parker, D.J., Allen, D.A., Benton, D.M., Fowles, P., McNeil, P.A., Tan, M., Beynon, T.D., 1997. Developments in particle tracking using the Birmingham Positron Camera. Nuclear Instruments and Methods in Physics Research A 392 (1–3), 421–426. Parker, D.J., Forster, R.N., Fowles, P., Takhar, P.S., 2002. Positron emission particle tracking using the new Birmingham positron camera. Nuclear Instruments and Methods in Physics Research A 477 (1–3), 540–545. Schaafsma, S.H., Marx, T., Hoffmann, A.C., 2006. Investigation of the particle flowpattern and segregation in tapered fluidized bed granulators. Chemical Engineering Science 61 (14), 4467–4475. Schlichting, H., Gersten, K., 1997. Grenzschicht–Theorie (Boundary Layer Theory). Springer, Berlin. Simons, S.J.R., 1995. Imaging techniques for fluidized bed systems: a review. Chemical Engineering Journal 56 (3), 83–93. Singh, R.K., Suryanarayana, A., Roy, G.K., 1992. Prediction of minimum velocity and minimum bed pressure drop for gas–solid fluidization in conical conduits. Canadian Journal of Chemical Engineering 70 (1), 185–189. Stein, M., Ding, Y.L., Seville, J.P.K., Parker, D.J., 2000. Solids motion in bubbling gas fluidised beds. Chemical Engineering Science 55 (22), 5291–5300. Stein, M., Ding, Y.L., Seville, J.P.K., 2002. Experimental verification of the scaling relationships for bubbling gas-fluidised beds using the PEPT technique. Chemical Engineering Science 57 (17), 3649–3658. Tanfara, H., Pugsley, T., Winters, C., 2002. Effect of particle size distribution on local voidage in a bench-scale conical fluidized bed dryer. Drying Technology 20 (6), 1273–1289. Tayebi, D., Svendsen, H.F., Jakobsen, H.A., Grislingås, A., 2001. Measurement techniques and data interpretations for validating CFD multi phase reactor models. Chemical Engineering Communications 186, 57–159. Teunou, E., Poncelet, D., 2002. Batch and continuous fluid bed coating – review and state of the art. Journal of Food Engineering 53 (4), 325–340. Vilstrup, P., 2001. Microencapsulation of Food Ingredients. Leatherhead Food International, Surrey. Watano, S., Yoshikawa, K., Miyanami, K., 1994. Evaluation of aqueous enteric coated granules prepared by moisture control method in tumbling fluidized bed process. Chemical and Pharmaceutical Bulletin 42 (3), 663–667. Werner, S.R.L., Jones, J.R., Paterson, A.H.J., Archer, R.H., Pearce, D.L., 2007. Airsuspension particle coating in the food industry: Part I – state of the art. Powder Technology 171 (1), 25–33. Werther, J., 1999. Measurement techniques in fluidized beds. Powder Technology 102 (1), 15–36. Wong, Y.S., Salleh, A.M., Ding, Y.L., Seville, J.P.K., Horio, M., 2001. Particle motion in relatively thin fluidised beds with immersed tubes. In: Proceedings of the International Congress for Particle Technology (Paper No. 248), PARTEC-2001, Nuremberg. Wurster, D.E., 1959. Air suspension technique of coating drug particles; a preliminary report. Journal of the American Pharmaceutical Association. American Pharmaceutical Association (Baltim) 48 (8), 451–454. Yang, Z., Parker, D.J., Fryer, P.J., Bakalis, S., Fan, X., 2006. Multiple-particle tracking – an improvement for positron particle tracking. Nuclear Instruments and Methods in Physics Research A 564 (1), 332–338. Yang, Z., Fan, X., Fryer, P.J., Parker, D.J., Bakalis, S., 2007. Improved multiple-particle tracking for studying flows in multiphase systems. AIChE Journal 53 (8), 1941– 1951. Yates, Y.G., Simons, S.J.R., 1994. Experimental methods in fluidization research. International Journal of Multiphase Flow 20 (Suppl. S), 297–330.