Bubble Characterization in a Fluidized Bed with Vertical Tubes

Article pubs.acs.org/IECR

Bubble Characterization in a Fluidized Bed with Vertical Tubes Martin Rüdisüli,† Tilman J. Schildhauer,*,† Serge M. A. Biollaz,† and J. Ruud van Ommen‡ †

Paul Scherrer Institut (PSI), General Energy Research Department, CH-5232 Villigen PSI, Switzerland Delft University of Technology, Product and Process Engineering, Julianalaan 136, 2628 BL Delft, The Netherlands

‡

ABSTRACT: Vertical tubes are commonly used in industrial fluidized beds as heat exchanger tubes. In this study, the influence of vertical tube banks on the axial bubble growth in a 145-mm inner diameter (ID) fluidized bed with Geldart B particles is experimentally investigated by means of pressure fluctuation measurement (PFM) and optical probing (OP). The employed tube bank configurations differ in their tube-to-tube spacing, tube diameter, and tube arrangement (square versus triangular). PFM and OP show that immersed vertical tube banks, irrespective of their configuration, significantly reduce the axial bubble growth. The bubble size reduction is even more significant for higher gas velocities (u0/umf > 4.6, where u0 is the superficial gas velocity and umf is the minimum fluidization velocity) as well as if the tube-to-tube spacing is narrow and the tube diameter is small. The tube arrangement (square versus triangular), in turn, does not show any significant differences. Depending on the gas velocity, the ratio of the bubble diameter to the tube diameter and the tube-to-tube spacing is changed. This change invokes a different bubble flow pattern between the vertical tubes and different bubble growth mechanism, which defines the effectiveness of vertical tubes to delay slugging and makes the fluidization smoother.

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INTRODUCTION Experimental studies with gas−solid fluidized beds at laboratory scale are usually conducted without internals. However, in industrial fluidized-bed applications, internals such as heatexchanger tubes and baffles are regularly employed. These immersed internals typically modify the gas−solid flow structure and thus may have many beneficial effects on the fluidization. First, internals can improve selectivity by preventing gas from backmixing.1 Second, bubble growth can be controlled better by reducing and homogenizing the bubble size distribution in the bed.2 This is achieved by preventing small bubbles from coalescing and enhance bubble splitting for large bubbles.3,4 Hence, slugging, bed height fluctuations, and particle elutriation is minimized, while local solids circulation is enhanced.3 Moreover, the bubble size reduction and the consequent redistribution of gas into the dense phase leads to an increased residence time of the gas in the bed and, therefore, a higher chemical conversion in the reactor.5 However, the complex hydrodynamics in bubbling fluidized beds (BFBs) with immersed internals are still difficult to describe. In particular, the interactions of the tubes with rising bubbles are difficult to understand. Consequently, fluidized-bed reactors with internals pose a substantial challenge and technological risk to plant designers and investors.6 In the past decades, the effect of internals in BFB has extensively been investigated experimentally and numerically (see, e.g., refs 6−10) and numerous comprehensive reviews have been published (see, e.g., refs 11−13). Historically, the first to investigate fluidized beds with internals were Glass and Harrison.14 They found that, for gas velocities slightly above umf, horizontal tubes improve bubble breakup and lead to “smoother” fluidization.15 The actual boom for research on internals began in the 1970s, with the emergence of pressurized fluidized-bed combustors (PFBCs), where the heat production per unit volume is high and internals are used as heat-exchanger tubes.16 Back then, primarily horizontal coils in a serpentine © 2012 American Chemical Society

configuration were used, but the high density of the coils resulted in poor heat exchange along the bed height, because of poor solids circulation and particle agglomeration. Therefore, vertical configurations were used with much better performance. Other early industrial applications of heat exchanger tubes in PFBCs (0.1−80 MW) can be found in the work of Zakkay et al.17 The effectiveness of internals is greatly dependent on their design. Typically, the design of internals in fluidized beds is categorized according to their type (as baffles, tubes, packings, inserted bodies, etc.)13 and their orientation in the bed (such as horizontal, vertical, or inclined). Generally, only the vertical and horizontal orientations are useful. Inclined orientation leads to excessive gas bypassing and decreased heat transfer, and it causes short-circuiting of gas along the undersides; therefore, it is not recommended.5,15 The choice whether to use horizontal or vertical tubes often depends on the aspect ratio H/D of the reactor. Because industrial fluidized-bed reactors are usually limited in height, horizontal tubes are employed more readily. Still, vertical tubes have many distinct advantages over horizontal tubes: simpler design, easier installation, and no upstream dead spots,5 as well as enhanced particle mobility, resistance to tube erosion, and higher heat loadings.17 In particular, reduced tube erosion is an obvious advantage of vertical tubes over horizontal tubes, since bubbles and particle jets smoothly and tangentially pass the tubes. Thus, tube erosion by mechanical wear for the same gas velocity is reduced by 50%.18 Also in the light of reactor scaleup, vertical tubes are employed to better control bubble growth.19 Received: Revised: Accepted: Published: 4748

September 30, 2011 February 23, 2012 February 26, 2012 March 6, 2012 dx.doi.org/10.1021/ie2022306 | Ind. Eng. Chem. Res. 2012, 51, 4748−4758

Industrial & Engineering Chemistry Research

Article

In industrial fluidized beds, vertical internals, such as vertical rods, vertical tubes, vertical helix springs, etc., are typically used as heat exchangers.4 Heat exchange with vertical tubes is slightly different from horizontal tubes, as explained in the work of Ozawa et al.20 A significant drawback of vertical tubes compared to freely bubbling fluidized beds is their reduction of axial and lateral solids mixing, such that particle segregation is promoted and proper fluidization becomes difficult due to “channeling”.2,5,19 In the literature, only limited research has been published on vertical tubes in fluidized bedsin particular at laboratory scale. While previous studies rather focused on the design and the general advantages of very particular configurations of vertical tubes in fluidized beds,5,10,17,19 in this paper, the axial distribution of the bubble growth is systematically investigated by means of experiments in a laboratory-scale fluidized bed (D = 0.145 m inner diameter (ID)), as a function of an array of different tube bank configurations. These tube bank configurations encompass a modified tube diameter, tube-to-tube spacing, and tube arrangement. The investigated bubble characteristics (viz, the bubble rise velocity (BRV) and the bubble size) are measured by means of pressure fluctuation measurement (PFM) and optical probing (OP).

Figure 1. U-shaped endings of the vertical tubes in the Sq/10/9 bank. (See Table 1 for a description of the naming conventions.) If tubes cannot be interconnected by a plastic “U”, a white plastic plug is inserted instead.

the glass column wall, the horizontal spacers are covered with duct tape. If square or triangular tube banks are inserted into a cylindrical column, it often happens that fairly large empty spaces occur between the wall of the glass column and the outermost rim of the tube bank. These large empty spaces may act as preferential pathways for rising gas bubbles and may corrupt the actual hydrodynamics in the bed. Therefore, these empty spaces must be avoided by filling them with smaller auxiliary tubes. The diameter of the auxiliary tubes is chosen such that the tube-to-tube spacing and other characteristics of the tube bank are maintained as well as possible. A list of the different tube banks used in this study, can be found in Table 1. For a more graphical representation of the different tube banks, refer to Figure 2. Along with the tube bank configuration, in terms of tube diameter, tube-to-tube spacing, and tube arrangement, as well as the number of principal and auxiliary tubes and the hydraulic diameter, the area reduction, and the bed height increase of an immersed tube bank, are tabulated in Table 1. For instance, with Tri/10/9 (i.e., a tube bank with a tube diameter of 10 mm, a tube-to-tube spacing of 9 mm, and a triangular tube arrangement (denoted as Tri)), a reduction of the cross-sectional area of −17.6% and an increase of the bed height of +21.4% is encountered. Most notably, Tri/10/9 and Sq/10/9 are the only tube banks where all configurational features, except for the tube arrangement, are identical. For hydrodynamical design purposes of fluidized beds with vertical tubes, the hydraulic diameter (Dhyd) of the bed, given as

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EXPERIMENTAL SECTION Apparatus and Particle Characterization. All experiments in this paper have been conducted in a glass column fluidized bed (Glas15). The diameter of the cylindrical glass column was 0.145 m, and the height of the cylindrical glass column was 0.930 m. In order to take pressure fluctuation measurements using a piezoelectric pressure transducer (Kistler, Type 7261) with a resolution of 1.5 Pa, several probe ports are positioned at the wall of the glass column, at heights of 0.027, 0.081, 0.135, 0.189, 0.243, 0.300, 0.350, 0.400, and 0.450 m. In addition, two ports for self-made optical-fiber probes are located at column heights of 0.243 and 0.45 m. For further information on the Glas15 and on the measurement techniques employed, refer to Rüdisüli et al.21,22 The bed material is a porous aluminum oxide (γ-Al2O3). The technical name of the powder is Puralox NWa-155 (referenced hereafter as NWA), and it is manufactured by Sasol Germany GmbH. The NWA mean particle size is 289 μm; its particle density is 1350 kg/m3, and the minimum fluidization velocity umf is ∼0.041 m/s. For further specifications on NWA, refer to Rüdisüli et al.23 Vertical Tube Banks. In order to investigate the influence of vertical tubes on the hydrodynamics (i.e., the bubble growth) in the Glas15, several vertical tube banks have been constructed. These tube banks vary in the diameter of the tubes, the tube-to-tube spacing, and the geometrical arrangement of the tubes (i.e., square versus equilateral triangle). All tube banks are made from hollow steel rods which represent mock-ups of heat-exchanger tubes. In industrial fluidized beds, heat exchanger tubes are often interconnected at their bottom to form a “U”-shape. This U-shape is mimicked by connecting a horizontal plastic tube between the hollow ends of two adjacent vertical steel rods (cf. Figure 1). Moreover, in order to minimize vibration and dislocation of the vertical tubes (in particular, those not connected with a “U”), the tube banks are stabilized and interconnected by horizontal wire spacers placed 0.4 m above the lower end of the tube (cf. Figure 2). As a precaution against scratches from these horizontal spacers on

Dhyd =

4A P

(1)

where A is the cross-sectional area available to the fluidized bed and P is the total wetted perimeter of the bed and tubes, is used instead of the column diameter D.19,24 Irrespective of the orientation and arrangement of the tubes, a rule of thumb by Grace and Harrison5 demands a tube-to-tube spacing at least 20−30 times greater than the mean particle size; otherwise, bridging or wedging of particles in the gaps between the tubes may occur with the consequence of dead regions and deteriorated particle mixing. With a mean particle size of NWA of 289 μm, the minimum tube-to-tube spacing is 8.7 mm, which is fulfilled for all tube bank configurations in this study. 4749

dx.doi.org/10.1021/ie2022306 | Ind. Eng. Chem. Res. 2012, 51, 4748−4758


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optical probes and/or they generate a pressure wave detected by the pressure sensors. The fact that small bubbles hit the optical probes less frequently than larger bubbles is addressed by the statistical data evaluation discussed in ref 21. Furthermore, if bubbles move upward between tubes or are wrapped around tubes, they still have to replace the dense phase, thus changing the static pressure head dynamically and, therefore, a pressure wave is generated.25 The bubble size from PFM is evaluated by means of the spectral data decomposition described by Schaaf et al.26 This method is based on the decomposition of the PFM time series into its coherent power spectral density (IOP) and incoherent power spectral density (PSD)that is, the global and local components of the pressure waves, respectivelywhile the integral of the latter yields a characteristic length scale proportional to the bubble size. In order to obtain the BRV and bubble size from OP, a bubble linking algorithm described in Rüdisüli et al.22 is used. This bubble linking algorithm based on least-squares regression techniques evaluates bubble characteristics by measuring the response of a bubble pierced by two reflective optical fiber probes spaced at 1 cm. For more information on the technical details, the installation, and the application of PFM and OP refer to Rüdisüli et al.23 and Rüdisüli et al.,21,22 respectively. As a third and complementary means to investigate bubble characteristics, visual observations also are made in the glass column fluidized bed (Glas15). Because of geometrical constraints with triangular tube banks, OP can only be applied to some square internals, viz, Sq/15/9, Sq/15/18, and Sq/20/9, where the optical probes can be inserted between the tubes. By default, optical probe measurements are taken with the probe tip at 2 cm from the center of the bed (∼D/3).22 Experiments are conducted at relative gas velocities (u0/umf, where u0 is the superficial gas velocity and umf is the minimum fluidization velocity) of 2.3, 4.5, and 6.8, such that experiments with internals can hydrodynamically be compared to experiments without internals in Rüdisüli et al.22 Because of different area reductions of the tube banks (cf. Table 1), the actual gas flow rate with internals is adjusted to the gas velocity u0/umf. This means that the gas flow rate must be decreased by the corresponding area reduction of the used tube bank. For example, if the bed without internals is operated at 100 NL/min, the bed with internals (e.g., Sq/15/9) has been decreased by −22.6% to 77.4 NL/min in order to have the same u0/umf. The default distance between the distributor plate and the lower end of the tube bank is 10 mm. In order to see the influence of this distance on the bubble size, experiments with Sq/15/9 (i.e., the reference tube bank) are repeated with the tube bank raised by 60 mm and 120 mm. Moreover, also the influence of the azimuthal angle of the tube banks on the PFM

Figure 2. Employed vertical tube banks. Principal tubes are represented by dark blue dots. Smaller auxiliary tubes are displayed as light blue dots. “U”-shaped connectors at the bottom of the tubes are shown as green bars and horizontal support spacers are given as black lines. The default measurement location for optical probing (OP) is marked with a red cross (×). The reference case (a) Sq/15/9 also shows the different azimuthal angles (0°, 45°, and 90°) of the PFM sensors. (See Table 1 for a description of the naming conventions.)

Experiments. In order to measure bubble characteristics in the fluidized bed with vertical tubesviz, the bubble rise velocity (BRV) and the bubble sizepressure fluctuation measurement (PFM) and optical probing (OP) are employed. Both methods are suitable as long as bubbles are pierced by the

Table 1. Physical Properties of the Tube Bank Configurations Used in This Study

a

name

tube arrangement

tube diameter [mm]

tube spacing [mm]

principal tubes

auxiliary tubes

hydraulic diametera [mm]

area reduction [%]

height increase [%]

Sq/15/9 Sq/15/18 Sq/20/9 Sq/10/9 Tri/10/9 Tri/15/9

square square square square triangular triangular

15 15 20 10 10 15

9 18 9 9 9 9

16 12 12 37 37 24

8 (12 mm) 4 (5 mm) 4 (9 mm)

33.8 52.8 37.8 33.6 33.6 29.4

−22.6 −13.3 −24.4 −17.6 −17.6 −26.6

29.2 15.4 32.2 21.4 21.4 36.3

2 (10 mm)

The hydraulic diameter is calculated using eq 1. 4750



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and the bubble size is investigated by azimuthally shifting Sq/ 15/9 by 45° and 90° (cf. Figure 2a).

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RESULTS Pressure Fluctuation Measurement. Spectral Data Decomposition. Figure 3 illustrates PFM time series without internals

Figure 4. Spectral data decomposition of the pressure fluctuation time series (a) without internals and (b) with Sq/15/9 internals (cf. Figures 3a and 3b, respectively) at u0/umf = 4.5.

This unimodal frequency distribution of the PSD is in contrast to the PSD of PFM at the same gas velocity u0/umf without internals in Figure 4a. Without internals at higher probe heights, a dominant and spiky peak is found at ∼1.5 Hz, which can be attributed to slugging in the upper portions of the bed. The absence of this spike in Figure 4b, along with a broader and more regular PSD, indicates that slugging can effectively be prevented by means of vertical tubes. The incoherent power spectral density (IOP) calculated by the spectral data decomposition method by Schaaf et al.26 with Sq/15/9 internals is displayed in Figure 5. As the PSD in Figure 4b, the corresponding IOP features a broad peak at ∼2.5 Hz. Since the IOP exclusively represents pressure fluctuations caused by bubble passage,26 the bubble size distribution in the bed can only be retrieved from the IOP. In this case, the bubble size distribution seems to be regular. Moreover, the total power in the IOP in Figure 5 is significantly lower than that of the PSD in Figure 4a. Therefore, coherent pressure fluctuations originating from global phenomena such

Figure 3. Pressure fluctuation time series at a relative gas velocity of u0/umf = 4.5 for a fluidized bed (a) with no internals and (b) with Sq/15/9 internals.

(Figure 3a) and with Sq/15/9 at u0/umf = 4.5 (Figure 3b). The time series 2 s in length are typical for the entire time series of 20 min. Apparently, the time series with internals (Sq/15/9) are more alike and the amplitudes of the fluctuations are lower than those without internals, irrespective of the height in the bed. This simple comparison of time series with and without internals already shows that fluidized beds with vertical tubes feature a “smoother” and more regular fluidization. The power spectral density (PSD) of the time series in Figures 3a and 3b at u0/umf = 4.5 are displayed in Figures 4a and 4b, respectively. The PSD of the time series with Sq/15/8 internals in Figure 4b shows a single (broad) peak at ∼2 Hz. 4751



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sizes at 90° are slightly larger than at the other two azimuthal probe angles. Another implication of the azimuthal angle is the relative position of the “U” at the bottom of the tubes to the PFM probe. This relative orientation of the “U” may have an influence on the rise pathways of bubbles. However, according to Figure 6, there is no effect of the orientation of the “U” at the bottom of the tubes on the measured bubble size. A potential reason why the azimuthal angle has no influence is the detection range of the PFM probes, which is on the order of 0.5 m for Geldart B particles.27 Since the column diameter of the “Glas15” (D = 0.145 m) is considerably smaller than this detection range, all rising bubbles in the column are measured in a holistic way, irrespective of their rise pathways. Distributor Distance. The influence of the distance from the tubes to the distributor plate on the bubble size is displayed in Figure 7 for Sq/15/9 and three (relative) gas velocities Figure 5. Incoherent power spectral density (IOP) of the pressure fluctuation time series with Sq/15/9 internals (cf. Figure 3b and Figure 4b) at u0/umf = 4.5, evaluated by means of the spectral data decomposition by Schaaf et al.26

as bubble formation, bubble coalescence, and bubble eruption26 constitute the bulk of the power in the pressure fluctuation signals. Azimuthal Angle. Depending on the azimuthal angle between the PFM probe and the vertical tubes (cf. Figure 2a), the immediate vicinity in front of the PFM probe tip is changed. If the azimuthal angle is 0° (default) or 90°, the two auxiliary tubes of Sq/15/9 are just next to the probe tip, and in direct and straight extension of the probe line, there are no tubes (this condition is considered to be an open throughfare channel). On the other hand side, at an azimuthal angle of 45°, there is more empty space just next to the probe, yet there is no open throughfare channel in direct extension of the probe. Figure 6 displays the influence of the azimuthal probe angle on the measured bubble size. The bubble size is represented as

Figure 7. Bubble size relative to the distance between the tube bank and the distributor plate with Sq/15/9 internals.

(u0/umf = 2.3, 4.5, and 6.8). The investigated elevations of the tube bank are 0 mm (which is the default), 60 mm, and 120 mm. For u0/umf = 2.3, no difference is observed. At u0/umf = 4.5, a significant increase in the bubble size can only be noticed if the bottom of the tube bank is elevated by 120 mm. In contrast, at the highest gas velocity of u0/umf = 6.8, a substantial increase of the bubble size is already observed at an elevation of 60 mm. For an elevation of 120 mm at u0/umf = 6.8, the influence on the bubble size is even larger. The 120 mm curve and the default curve (0 mm) start to depart from each other at a height of ∼120 mm. Below that height, the two curves are more or less congruent. A reason why there is no influence for low gas velocity is that the gas velocity has a significant influence on the bubble size; that is, at low gas velocities, bubbles are significantly smaller than those formed at high gas velocities. Therefore, at low gas velocities, the influence and interactions of vertical tubes with bubbles in terms of bubble size reduction are also smaller than at high gas velocities (cf. the next subsection and the Discussion section). Influence of Vertical Tubes. The interactions of immersed vertical tube banks with rising gas bubbles is investigated at three relative gas velocities (u0/umf = 2.3, 4.5, and 6.8) and are displayed in Figure 8. Tube bank modifications of the reference

Figure 6. Bubble size, depending on the azimuthal angle between the PFM probe and the orientation of the Sq/15/9 internals.

the characteristic length scale of the bubble at a specific height.26 Apparently, there is no significant difference in the bubble size, depending on a variation of the azimuthal probe angle. For a given gas velocity, all bubble size curves are almost identical. Only at the lowest gas velocity of u0/umf = 2.3, bubble 4752



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For all gas velocities in Figure 8, the bubble size without internals (thick blue line) features a steady increase up to a height of ∼0.35 m. Above 0.35 m, the measured bubble size gradually levels off or even starts to decrease for higher gas velocities. This may indicate that a maximum bubble size is attained; however, a maximum attainable bubble size is rather found with Geldart A particles.24 Because the used NWA particles are Geldart B, this leveling off and decrease of the measured bubble size is generally an artifact of the bubble eruption at the bed surface.23,26 With internals, Figure 8 shows that all tube bank configurations feature a significant bubble size reduction compared to no internals. Generally, the most significant bubble size reductions are found for high gas velocities u0/umf = 4.5 and 6.8, which are more relevant for industrial fluidized beds. At u0/umf = 6.8, the observed bubble size reduction is even more pronounced than at u0/umf = 4.5. Moreover, the growth of the bubble size with internals is more linear and no decrease toward the bubble eruption zone is found. However, toward the eruption zone of the bubble, initially, similar bubble size curves start to depart from each other, while others approach again. These irregular patterns of the growth of the bubble size are also an artifact of the bubble eruption at the bed surface.26 For instance, without internals, bubbles grow faster and the fluidization regime is rather slugging than bubbling. Thus, the influence of the bubble eruption starts lower in the bed and is more dominant. Therefore, curves of no internals and curves with internals meet again. Below a height of 0.081 m, all bubble size curvesalso without internals and for high gas velocitiesare almost identical. An explanation is that, close to the distributor plate, bubbles are still small in size, and, therefore, no growth of the bubble size by coalescence and no interactions with the vertical tubes may have started yet. A comparison of the different tube bank configurations shows that a larger tube-to-tube spacing (Sq/15/18) and a larger tube diameter (Sq/20/9) is less effective to reduce the bubble size. On the other hand, a smaller tube diameter (Sq/ 10/9) results in more bubble size reduction. A smaller tube diameter also results in a larger number of tubes per crosssectional area which may be the principal reason for enhanced bubble size reduction. Enhanced bubble size reduction can also be found for triangular tube arrangements (i.e., Tri/10/9 and Tri/15/9). For a more conclusive discussion of the influence of different tube configurations on the bubble size reduction, refer to the Discussion section. Optical Probing. Figure 9 shows the normalized histograms of the bubble rise velocity (BRV) with internals (Sq/15/ 9, Sq/15/18, and Sq/20/9) and without internals measured by optical probes. The optical probe heights are 0.243 and 0.450 m, while the gas velocities are u0/umf = 2.3, 4.6, and 6.8. Along with the BRV distributions, their means also are displayed. For u0/umf = 2.3, a decrease of the BRV, compared to no internals, is obtained for all internals and at both probe heights. Contrarily, for u0/umf = 4.6, the BRV is increased for all internals and at both probe heights. Eventually at u0/umf = 6.8, a decrease of the BRV is observed at 0.243 m and an increase at 0.450 m. This shows that, depending on the gas velocity and the height in the bed, no clear trend about the effect of internals on the BRV can be given. Reasons for an increased BRV with internals are given in the literature: Grace and Harrison5 reported that bubbles enclosing vertical tubes and bubble columns elongate considerably in the vertical direction and rise up more rapidly. Similar phenomena have also been witnessed with bubbles between vertical tubes in

Figure 8. Summary of the characteristic length scale obtained from PFM for different tube bank configurations at relative gas velocities of (a) u0/umf = 2.3, (b) u0/umf = 4.5, and (c) u0/umf = 6.8.

tube bank Sq/15/9 (cf. Table 1) include a wider tube-totube spacing (Sq/15/18), a larger and smaller tube diameter (Sq/20/9 and Sq/10/9), and a modified tube arrangement (Tri/15/9), as well as both effects (tube diameter and arrangement) at the same time (Tri/10/9). 4753



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Regarding the influence of vertical tubes on the bubble size distribution, Figure 10 shows that the mean bubble size is

Figure 9. Normalized histograms of the bubble rise velocity (BRV) distribution. Tube bank configurations are Sq/15/9 (reference), Sq/15/18 (wider tube-to-tube spacing), and Sq/20/9 (larger tube diameter), as well as no internals. Optical probe heights are (a) 0.243 m and (b) 0.450 m.

Figure 10. Kernel smoothed density of the bubble size distribution. Tube bank configurations are Sq/15/9 (reference), Sq/15/18 (wider tube-to-tube spacing), and Sq/20/9 (larger tube diameter), as well as no internals. Optical probe heights are (a) 0.243 m and (b) 0.450 m.

gas−solid fluidized beds with large particles: Bubbles tend to elongate while they travel upward in a faster and less oblique way.5 Gallucci et al.28 showed that the BRV increases with the number of vertical tubes, because of the preferential pathways along the tube walls. Yates et al.29 indicated that, particularly, tubes with roughened surfaces feature a higher BRV than smoothly polished ones. In this manner, the BRV can become twice as fast as with bubbles of the same volume in the absence of tubes. Although literature indicates that the BRV rather increases between vertical tubes, the exact relationships must be examined more thoroughly.

substantially decreased with internals, compared to without internals. Generally, the bubble size reduction is more significant for higher gas velocities. This seems to be evident, since vertical tubes avoid slugging, which is particularly prominent for high gas velocities. Furthermore, not only is the mean bubble size reduced, but the shape of the bubble size distribution also is altered. For all gas velocities, the bubble size distribution with internals is more right-skewed and narrower. The right-skewness follows from 4754



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Figure 11. Summary of the percentage bubble size reduction with optical probing (OP) and pressure fluctuation measurement (PFM) at a height of z = 0.243 m.

methods yield similar trends and similar proportions in their bubble size reduction for the reference internals (Sq/15/9) and for a wider tube-to-tube spacing (Sq/15/18). In contrast, for a larger tube diameter (Sq/20/9), the same trends between OP and PFM can only be seen for u0/umf = 2.3, while for higher gas velocities, OP indicates that Sq/20/9 yields less bubble size reduction than Sq/15/18. This is opposed to PFM, where the bubble size reduction of the two configurations is just the opposite. A potential reason for this discrepancy is that optical probes give very local information on the bubble, depending on the surrounding tubes in Figure 2, whereas the PFM gives an “averaged” value over the full horizontal cross section. Another finding from Figure 11 is that, at u0/umf = 2.3, Tri/ 10/9 shows the least bubble size reduction of all tube banks, while the same configuration shows the most-efficient bubble size reduction for higher gas velocities. Generally, as already shown in the Results section for both PFM and OP, the bubble size reduction for u0/umf = 4.5 and 6.8 are very comparable; however, the trends in the bubble size reduction at u0/umf = 2.3 are quite different. Consequently, there must be a significant influence of the gas velocity on the efficiency of the vertical tubes on the bubble size reduction. In other word, there may be a transition between u0/umf = 2.3 and 4.5, where the bubble size reduction characteristics and the interactions of the bubble with the vertical tubes change quite substantially. In order to explain this transition, the mechanism of bubble size reduction and the effects of the different configurations viz, tube diameter, tube-to-tube spacing, and tube arrangementmust be examined separately and compared to the literature. Vertical tubes act on the bubble size in two ways: (i) they avoid coalescence and (ii) they enhance bubble splitting. While the avoidance of coalescence is rather attributed to the tube diameter,5 the enhancement of bubble splitting is attributed to the tube-to-tube spacing.29 In this respect, special attention is paid to the ratio of the bubble diameter to the tube diameter and the ratio of the bubble diameter to the tubeto-tube spacing. Regarding the influence of the tube arrangement (triangular versus square), there is no significant difference in the achieved bubble size reduction in Figure 11. Therefore, the

the fact that tubes act as obstacles and, consequently, bubbles cannot grow to large sizes by coalescence so easily. Despite the general bubble size reduction, bubbles still reach sizes significantly larger than the tube-to-tube spacing and the hydraulic diameter of Sq/15/9, Sq/15/18, and Sq/20/9 (cf. Table 1). This is in contradiction to the theory of Ozawa et al.,20 who proposed to estimate the bubble size in a fluidized bed with vertical tubes by taking the bubble size correlation of Mori and Wen30 and just replacing the cross-sectional area by the square of the tube-to-tube spacing in the calculation of the maximum attainable bubble size.

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DISCUSSION Irrespective of the tube configuration and the measurement methodpressure fluctuation measurement (PFM) and optical probing (OP)there is a substantial influence of vertical tubes on the bubble size in a fluidized bed. For all experiments in this paper, for a given relative gas velocity u0/umf, a significant reduction of the bubble size has been observed. Figure 11 summarizes the percental bubble size reduction at a height of 0.243 m for all vertical tube bank configurations, all gas velocities (u0/umf = 2.3, 4.5, 6.8) and both measurement techniques. A probe height of 0.243 m is taken as reference, since it is the most reliable height for PFM, with a minimum of signal distortion by bubble eruption and bubble formation.23 Figure 11 shows that, with OP, less bubble size reduction is measured than with PFM. This discrepancy involves the different physical principles of the two measurement methods. OP is a direct method, where the size of ellipsoidal bubbles is approximated by the vertical chord length of pierced bubbles, as statistically demonstrated by Rüdisüli et al.21 Alternatively, PFM is an indirect method, where the bubble size is estimated from the integral of incoherent pressure fluctuations.26 Therefore, PFM is not dependent on the bubble shape and only provides an integral characteristic length scale proportional to the absolute volume-equivalent bubble size.26 Since vertical tubes typically elongate the shape of a bubble,31 the pierced chord length of OP intrinsically yields less bubble size reduction than the integral and bubble-shape-independent characteristic length scale of PFM. Nonetheless, both measurement 4755



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Eventually, also from visual observations of the fluidization with vertical tubes in this study, a “smoother” fluidization can be observed. This “smoother” fluidization is characterized by less piston-like oscillations of the bed mass, a reduction of slugging, more gentle bubble eruptions at the bed surface, and smaller and more homogeneous bubbles. This evidence of a “smoother” fluidization can be summarized as the fluidized bed adopting bubblingnot sluggingfluidization. Since, for most industrial applications, the avoidance of slugging is preferred, vertical internals help improve the efficiency and the overall profitability of bubbling fluidized-bed reactors. On these grounds, the optimal choice of a tube bank configuration should be adjusted to the typical gas velocity at which the unit is operated.

arrangement of tubes in squares or equilateral triangles does not have a significant influence on the bubble size. The only exception of this rule is found for u0/umf = 2.3 and Tri/10/9, where the bubble size reduction is significantly smaller than its square counterpart (Sq/10/9), which has exactly the same physical properties as Tri/10/9 (cf. Table 1). A physical reason for this inconsistent discrepancy is not known. As previously mentioned, the ratio of the bubble diameter to the tube diameter (dB/dtube) is of utmost importance to understanding the influence of vertical tubes on the bubble size. According to Grace and Harrison,5 if the ratio dB/dtube is small (i.e., the tube diameter is larger or similar to the bubble diameter), bubbles move and coalesce undisturbedly between the vertical tubes.32 A graphical illustration of this situation can be found in Figure 12a. Because bubbles are typically small for

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CONCLUSION The influence of vertical tube banks on the axial bubble rise velocity (BRV) and the bubble size in a fluidized bed with Geldart B particles has been investigated for (relative) gas velocities of u0/umf = 2.3, 4.5, and 6.8, by means of pressure fluctuation measurement (PFM) and optical probing (OP). Employed vertical tube bank configurations encompass the variation of the tube-to-tube spacing, the tube diameter, and the tube arrangement (square versus triangular). Time series analysis of pressure fluctuations by means of spectral data decomposition26 shows that, for a given gas velocity u0/umf, vertical tubes significantly influence the hydrodynamics of the fluidized bed. In particular, slugging fluidization is inhibited and smoother fluidization is promoted. These observations from time series analysis coincide with visual observations in the glass column. PFM of the bubble size with the tube banks shifted by 45° and 90° show that there is no influence of the azimuthal probe angle. However, it must be noted that this may only hold for fluidized beds, where the column diameter is reasonably smaller than the detection range of the pressure sensors (i.e., ∼0.5 m for Geldart B particles27). Furthermore, depending on the gas velocity, an influence of the distance between the tube bank and the distributor plate on the bubble size is shown from PFM. For increased gas velocities, a wider distance results in a significantly faster axial growth of the bubble size. On the other hand, if the gas velocity is low (u0/umf = 2.3), no influence of this distance is measured. Generally, for a given gas velocity u0/umf, the axial growth of the bubble size measured by both methodsPFM and OP has shown a significant reduction of the bubble size for all tube bank configurations. This is particularly true for high gas velocities and in the upper portions of the bed, where slugging typically occurs without internals. The best results, with respect to an efficient bubble size reduction, are obtained for a narrow tube-to-tube spacing and small tube diameter. Small tube diameter, in turn, allow for a large number of tubes per crosssectional area. In contrast, no significant difference is measured between square and triangular tube arrangements. The mechanism of bubble size reduction is clearly dependent on the gas velocity and the consequent ratio of the bubble diameter to the tube diameter and tube-to-tube spacing, respectively. If the tube diameter is larger than the bubble diameter, the bed cross section is divided into several parallel sectors, where bubbles can move and coalesce undisturbedly. Therefore, a more efficient bubble size reduction is achieved if the tube diameter is smaller than the bubble diameterthis is typically the case if the gas velocity is highand bubbles may

Figure 12. Interaction of bubbles with vertical tubes depending on the ratio of the bubble diameter dB to the tube diameter dtube for (a) dB < dtube and (b) dB > dtube.

low gas velocities, this may explain the discrepancy in the bubble size reduction between low and high gas velocities in Figure 11. On the other hand, if the ratio dB/dtube is greater than 5 (i.e., the bubble is significantly larger than the tube diameter), several bubbles will enclose the vertical tubes and rise up along their walls, as if they were in an elevator (cf. Figure 12b).5 As they rise upward and cling to the tube, they are elongated, accelerated, and stabilized.31 In this manner, coalescence with other bubbles is hampered and the growth of the bubble size and potential slugging is efficiently delayed.5 In summary, if the tube diameter is larger than the bubble diameter, the vertical tubes act like outer walls, dividing the cross section into parallel chambers (called sectors), whereas if the tube diameter is smaller than the bubble diameter, bubble enclosure and bubble size reduction takes place. Regarding the influence of the tube-to-tube spacing, a comparison between Sq/15/9 and Sq/15/18 in Figure 11 shows that a wide tube-to-tube spacing (Sq/15/18) is less efficient in the bubble size reduction. Yates et al.29 claimed that the influence of two vertical tubes on the bubble size starts as they are spaced slightly further apart than the bubble diameter. At a certain distance from each other, vertical tubes start to break up the bubble and the two newly formed bubbles rise up along the tube walls. With three vertical tubes in a triangular arrangement and at a particular separation, bubbles are split into three daughter bubbles rising up along each tube.29 With the aid of neutron radiography, Ozawa et al.33 found that if the tube-to-tube distance becomes too wide, the effect on bubble size reduction becomes small. In addition, for very small bubbles, the tube-to-tube distance has a weak influence. 4756



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immersed tubes in bubbling fluidized beds. Powder Technol. 2004, 145, 88−105. (7) Olowson, P. A. Influence of pressure and fluidization velocity on the hydrodynamics of a fluidized bed containing horizontal tubes. Chem. Eng. Sci. 1994, 49, 2437−2446. (8) Hull, A. S.; Chen, Z.; Fritz, J. W.; Agarwal, P. K. Influence of horizontal tube banks on the bubbling and solids mixing behavior of fluidized beds. Available via the Internet at http://lib.kier.re.kr/ balpyo/15icfbc/99-0152.PDF, 1999. (9) Gustavsson, M.; Almstedt, A. E. Numerical simulation of fluid dynamics in fluidized beds with horizontal heat exchanger tubes. Chem. Eng. Sci. 2000, 55, 857−866. (10) Furui, S.; Umekawa, H.; Hayashi, K.; Ozawa, M.; Takenaka, N. Bubble behavior in vertical tube banks installed in a fluidized bed. Heat Transfer−Asian Res. 2003, 32, 727−739. (11) Harrison, D.; Grace, J. R. In Fluidization; Davidson, J., Harrison, D., Eds.; Academic Press: London and New York, 1971; Chapter 13, pp 599−626. (12) Sitnai, O.; Whitehead, A. B. In Fluidization, 2nd ed.; Davidson, J., Clift, R., Harrison, D., Eds.; Academic Press: London, 1985; Chapter 14, pp 473−493. (13) Yang, W. C. Handbook of Fluidization and Fluid−Particle Systems, 1st ed.; Marcel Dekker: New York, 2003. (14) Glass, D. H.; Harrison, D. Flow patterns near a solid obstacle in a fluidized bed. Chem. Eng. Sci. 1964, 19, 1001−1002. (15) Gogolek, P. E. G.; Grace, J. R. Fundamental hydrodynamics related to pressurized fluidized bed combustion. Prog. Energy Combust. Sci. 1995, 21, 419−451. (16) Gustavsson, M.; Almstedt, A. E. Two-fluid modelling of coolingtube erosion in a fluidized bed. Chem. Eng. Sci. 2000, 55, 867−879. (17) Zakkay, V.; Sellakumar, K. M.; Radhadrishnan, R.; McClung, J. D. Vertical heat exchanger for high pressure fluidized bed coal combustors. Energy Progress 1986, 6, 248−253. (18) Li, C.; Zakkay, V. Hydrodynamics and erosion modeling of fluidized bed combustors. J. Fluids Eng. 1994, 116, 746−755. (19) Volk, W.; Johnson, C. A.; Stotler, H. H. Effect of reactor internals on quality of fluidization. Chem. Eng. Prog. 1962, 58, 44−47. (20) Ozawa, M.; Umekawa, H.; Furui, S.; Hayashi, K.; Takenaka, N. Bubble behavior and void fraction fluctuation in vertical tube banks immersed in a gas-solid fluidized-bed model. Exp. Therm. Fluid Sci. 2002, 26, 643−652. (21) Rüdisüli, M.; Schildhauer, T. J.; Biollaz, S. M. A.; van Ommen, J. R. Monte Carlo simulation of the bubble size distribution in a fluidized bed with intrusive probes. Int. J. Multiphase Flow 2011, submitted. (22) Rüdisüli, M.; Schildhauer, T. J.; Biollaz, S. M. A.; van Ommen, J. R. Bubble characterization in a fluidized bed by means of optical probes. Int. J. Multiphase Flow 2012, 41, 56−67. (23) Rüdisüli, M.; Schildhauer, T. J.; Biollaz, S. M. A.; Wokaun, A.; van Ommen, J. R. Comparison of bubble growth obtained from pressure fluctuation measurements to optical probing and literature correlations. Chem. Eng. Sci. 2012, DOI: 10.1016.j.ces.2012.01.045. (24) Kunii, D.; Levenspiel, O. Fluidization Engineering, 2nd ed.; Butterworth−Heinemann: Boston, 1991. (25) van der Schaaf, J.; Schouten, J.; van den Bleek, C. Origin, propagation and attenuation of pressure waves in gas-solid fluidized beds. Powder Technol. 1998, 95, 220−233. (26) van der Schaaf, J.; Schouten, J. C.; Johnsson, F.; van den Bleek, C. M. Nonintrusive determination of bubble and slug length scales in fluidized beds by decomposition of the power spectral density of pressure time series. Int. J. Multiphase Flow 2002, 28, 865−880. (27) van Ommen, J. R.; van der Schaaf, J.; Schouten, J. C.; van Wachem, B. G. M.; Coppens, M.-O.; van den Bleek, C. M. Optimal placement of probes for dynamic pressure measurements in large scale fluidized beds. Powder Technol. 2004, 139, 264−276. (28) Gallucci, K.; Jand, N.; Foscolo, P. U.; Santini, M. Cold model characterisation of a fluidised bed catalytic reactor by means of instantaneous pressure measurements. Chem. Eng. J. 2002, 87, 61−71.

envelope the vertical tube as they rise upward along its outer wall, as if they were in an elevator. In this case, rising bubbles are efficiently separated from other bubbles and coalescence is inhibited.

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AUTHOR INFORMATION

Corresponding Author

*E-mail: [email protected]. Notes

The authors declare no competing financial interest.

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ACKNOWLEDGMENTS The authors would like to gratefully acknowledge “Verband Schweizerische Gasindustrie” (VSG/ASIG) for their funding of this project.

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NOMENCLATURE

Latin Symbols

A = cross-sectional area [m2] a = annular area [%] D = column diameter [m] H = column height [m] i = incremental number [−] j = incremental number [−] n = number of bubbles analyzed [−] P = wetted perimeter [m] p = probe position [m] w = number of bubbles analyzed [%] z = bed/probe height [m] dB = bubble size/diameter [m] umf = minimum fluidization velocity [m s−1] u0 = superficial gas velocity [m s−1] Subscripts and Superscripts

hyd = hydraulic tube = vertical tube Abbreviations

BFB = bubbling fluidized bed COP = coherent output-PSD ID = inner diameter IOP = incoherent output-PSD NWA = Puralox NWa-155 particles OP = optical probing PFBC = pressurized fluidized-bed combustor PFM = pressure fluctuation measurement PSD = power spectral density

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REFERENCES

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(29) Yates, J. G.; Cheesman, D. J.; Mashingaidze, T. A.; Howe, C.; Jefferis, G. The effect of vertical rods on bubbles in gas fluidized beds. In Fluidization; Kunii, D., Toei, R., Eds.; 1984. (30) Mori, S.; Wen, C. Y. Estimation of bubble diameter in gaseous fluidized beds. AIChE J. 1975, 21, 109−115. (31) Rowe, P. N.; Masson, H. Interaction of bubbles with probes in gas fluidised beds. Trans. Inst. Chem. Eng. 1981, 59, 177−185. (32) Rowe, P. N.; Everett, D. J. Fluidised Bed Bubbles Viewed by X-rays, Part I. Experimental Details and the Interaction of Bubbles with Solid Surfaces. Trans. Inst. Chem. Eng. 1972, 50a, 42−48. (33) Ozawa, M.; Umekawa, H.; Furui, S.; Hayashi, K.; Takenaka, N. Quantitative flow visualization of fluidized-bed heat exchanger by neutron radiography. Appl. Radiat. Isot. 2004, 61, 715−724.

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Bubble Characterization in a Fluidized Bed with Vertical Tubes

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