KONA Powder and Particle Journal No. 41 (2024) 140–150/https://doi.org/10.14356/kona.2024009

Review Paper

Applying Bubbling Fluidized-Bed Reactors for Strongly
Exothermic Reactions: Focus on Methanation †Philipp Riechmann and Tilman J. Schildhauer *Paul Scherrer Institute (PSI), Switzerland

1. Introduction

Many processes in the chemical industry and those converting chemical energy carriers are either strongly endoor exothermic. Most hydrogen-producing reactions areendothermic e.g. steam reforming, gasification, reverse
water gas shift, dehydrogenations, and methane dehydroaromatization to benzene. In consequence, many reactions that consume hydrogen are exothermic, e.g.
hydrogenation of double bonds or aromatic systems,
water–gas shift reaction, and important fuel syntheses suchas methanation, Fischer–Tropsch and methanol synthesis.
While endothermic reactions lead to the formation of acold spot where the further conversion of the reactant completely depends on further heat supply, strongly exothermic
reactions inherently bring the risk of thermal reactor runaway with them leading to a significant increase in temperature, formation of hot spots, often low selectivity, andpotential catalyst deactivation. For such reaction systems,
either highly temperature-stable catalysts or reactors withexceptional heat removal performance are required.
†*

Received 21 May 2022; Accepted 6 September 2022
J-STAGE Advance published online 30 June 2023
Corresponding author: Tilman J. Schildhauer;
Add: Forschungsstrasse 111, 5232 Villigen PSI, Switzerland
E-mail: tilman.schildhauer@psi.ch
TEL: +41-56-310-2706

140

Bubble Holdup εb [-]α [W/m²K]

300.0
0.20
Strongly exothermic reactions inherently pose the risk of thermal
α R0
0.18
α R3
reactor runaway, which may lead to very high increase in temperature,
250.0
Bubble Holdup R0
0.16
hot spots and potential catalyst deactivation. For such reaction systems,
Bubble Holdup R3
0.14
200.0
reactors with excellent heat removal performance are needed. In the
0.12case of methanation of CO/CO2-rich gases, full conversion is not
0.10
150.0
possible in a single adiabatic reactor due to the equilibrium limitation,
0.08
R0
R3
and in large-scale plants, e.g. coal-to-synthetic natural gas (SNG)
100.0
0.06
plants, series of four and more reactors with intercooling have been
0.04
50.0
realized. To allow for complete conversion within one reactor, the
0.02
potential of bubbling fluidized bed (BFB) reactors with immersed heat
0.00
0.0
1.5
2
3
4
exchangers was investigated in the US and Germany from the 1960s to
u/umf
the 1980s. A Swiss consortium started to expand the concept to smalland medium-scale plants to allow the production of renewable methanefrom decentral renewable sources such as wood gasification and biogas.
During their tests, it could be shown that the catalyst particle movement does not only allow for optimal heat removal—close toisothermal operation and thus little catalyst sintering—but that the catalyst particle movement over the height of the reactor with different
concentration zones favors the chemical catalyst stability. This contribution will review the fluid-dynamic studies for BFB reactors withimmersed heat exchangers in the last decades comprising studies with pressure fluctuation probes, optical probes, X-ray tomography
studies, and particle attrition studies.
Keywords: Bubbling Fluidized Bed (BFB) reactors, methanation, immersed heat exchanger, chemical catalyst stability, heat transfer

The methanation reaction is a typical example of ahighly exothermic and equilibrium-limited reaction. The
equilibrium is generally favored by lower temperatures ofabout 100 °C, where the conversion rate is close to 100 %(Witte, 2018). However, due to the low activity of the typically used catalysts at temperatures below 200–300 °C,
catalytic methanation reactors are usually operated above
this temperature level, at temperatures between 200 °C and
500 °C (Götz et al., 2016; Mills and Steffgen, 1974). The
typical pressure range is between 3 and 30 × 105 Pa withhigher pressures leading to slightly increased methane
yields (Götz et al., 2016). While ruthenium is the most active catalyst, it is seldom used due to its high cost. Nickel ishighly active, very selective and relatively cheap. It istherefore the preeminent catalyst material for catalytic
methanation processes (Mills and Steffgen, 1974; Ghaib etal., 2016).

2. Reactor types for methanation

Commonly used reactor types for the catalytic methanation process are series of adiabatic fixed-bed reactors,
cooled fixed-bed reactors, fluidized-bed reactors, slurry reactors, and metal monolith reactors (Schildhauer andBiollaz, 2015; Schildhauer, 2016). Fig. 1 shows illustrations of some of these reactor types. The requirements forthe reactor primarily depend on the composition and flow

Copyright © 2024 The Authors. Published by Hosokawa Powder Technology Foundation. This is an open
access article under the CC BY license (https://creativecommons.org/licenses/by/4.0/).

Philipp Riechmann et al. / KONA Powder and Particle Journal No. 41 (2024) 140–150

Review Paper

Fig. 1 Schematic representation of typical reactor types for methanation. (a): Multi-tubular cooled fixed bed, (b): Fluidized bed, (c): Stirred slurry
bubble column.

rates of the gas sources, as they determine the requirements
for the removal of the reaction heat and the prevention of
catalyst deactivation (Kopyscinski et al., 2010; Schildhauer,
2016; Schildhauer and Biollaz, 2015).
Using fixed-bed methanation reactors is an establishedmethod for gas cleaning in the ammonia synthesis process
chain. In this application, relatively small concentrations ofcarbon oxides have to be eliminated from the hydrogenrich feed streams to avoid catalyst deactivation. No additional cooling is required in this case due to the large ratio
between the inert gas and reactant gas volumes and the resulting relatively large heat capacity of the gas stream. For
the production of Synthetic Natural Gas (SNG), however,concentrations of inert gases are very low and cooling becomes necessary. This can be achieved through intermittent
cooling of the gas stream between multiple fixed-bed reactor stages. Additionally, gas can be recirculated for cooling.
Adiabatic reactors require robust catalyst materials that canwithstand a broad temperature range without sintering or
cracking (Bartholomew, 2001). For the production of SNG
from coal gasification gas, this reactor type is wellestablished, and plants at the GW-scale have been realized(Schildhauer, 2016). It is also possible to directly cool themulti-tubular fixed-bed reactors by having cooling media
flow around the catalyst-filled tubes.
Metal monolith reactors are designed to overcome someof the drawbacks of fixed-bed reactors. They are characterized by their high specific surface area, which allows relatively compact reactor designs. Their metal framework,
usually based on aluminum or copper, increases the radial
heat transfer, while the pressure drop is reduced significantly in comparison to fixed-bed reactors (Schildhauer,
2016; Schlereth et al., 2015). In this way, heat transfer coefficients of more than 1000 W/m2 K can be reached(Boger and Heibel, 2015; Schildhauer, 2016). However, thedeposition of the catalyst material on the structured framework is technically complex, especially when the catalyst

is deactivated and the reactor has to be coated again withthe catalyst.
In three-phase reactors, also known as slurry bubble
column reactors, the catalyst is suspended in an inert liquidby the reactant gas flow. On the one hand, the liquid phaseis beneficial as it significantly facilitates the heat removal,
leading to almost isothermal conditions within the reactor
and allowing a dynamic plant operation, where the reactor
temperature can be kept stable due to the thermal inertia ofthe liquid phase. On the other hand, the liquid phase induces an additional mass transfer resistance and needs tofulfill several specifications such as high thermal stability,
chemical inertness, low evaporation rates, low viscosity,and surface tension (Götz et al., 2013; Schildhauer, 2016).
Gas-solid fluidized-bed reactors are characterized bynearly isothermal conditions resulting from intensive particle circulation. Bubbles play a crucial role in catalytic
bubbling fluidized bed (BFB) reactors as they significantly
influence key properties of the fluidized bed, such as theheat and mass transfer, the mixing of gas and solids as wellas elutriation of bed material. Rising bubbles transportparticles upwards in their wake, which causes an intensive
particle motion resulting in a uniform temperature profileof the fluidized bed. Consequently, the heat transfer between the bed and heat exchanger surfaces is intensified
through the steady replacement of the particles. Due to theeven heat distribution throughout the reactor, the entire
immersed heat exchanger surface can be used at the sametime, which is the most important difference from fixedbed reactors. In this way, local temperature peaks and any
associated thermal runaway of reactions can be effectively
avoided. Therefore, BFBs with heat exchangers are well
suited for highly exothermic reactions such as catalytic
methanation. The properties of the bubble phase, including
intensity of particle circulation, are dependent on the ratio
between the superficial gas velocity u and the minimum
fluidization velocity umf, which is referred to as the141Philipp Riechmann et al. / KONA Powder and Particle Journal No. 41 (2024) 140–150

fluidization number. By setting the u/umf, the reactor performance can be adjusted according to the requirements.
For instance, it is possible to recirculate part of the productgas stream or reduce the system pressure to increase thefluidization number and thus intensify the heat transfer ifnecessary.
A commonly stated disadvantage of fluidized-bed reactors is the attrition of the catalyst material resulting fromthe comparatively high mechanical stress exerted on theparticles through elastic collisions with other particles or
the reactor walls. Such attrition can lead to the deactivationand elutriation of the catalyst particles, which therefore
would have to be replaced periodically.
Additionally, the formation of bubbles may lead to a
bypass of reactants through the reactor within the bubbles
and thus to an incomplete conversion of the CO or CO2 in
the feed stream. Other frequently mentioned disadvantages
are the possible erosion of internals, such as heat exchanger
tubes, varying solids residence time distributions, and thefrequently complex hydrodynamics that make scaling-up
difficult (Götz et al., 2016; Yang, 2013; Yatesand Lettieri,2016). The choice of suitable bed materials and detailed
knowledge of the fluid dynamics of BFBs allows overcoming these challenges during scale-up.

3. Fluidized-bed reactors for methanation

Fluidized-bed methanation was developed since the late
1950s to convert synthesis gas from coal gasification(Kopyscinski et al., 2010). The Bureau of Mines developed
a relatively slender reactor with a 2.5 cm diameter, operated with Geldart A type catalyst particles at 20 bar andaround 400 °C. Three feed injections were used to avoid
the strong axial temperature gradients observed when fed
only from the bottom. In 1963, the Bituminous Coal Research Inc. continued in the Bigas project with a 150 mm
diameter reactor, about 2 m bed height and two injection
points that were cooled with immersed heat exchanger
tubes. At temperatures from 430 to 530 °C, pressures from
6.9 to 8.7 × 105 Pa, it was operated as a BFB for several
1000 hours (Kopyscinski et al., 2010).
Within the Comflux process, the technology was scaled
up to the 20 MW scale in the early 1980s by Didier
Engineering GmbH and Thyssengas GmbH, supported by
the Technical University of Karlsruhe, and tested up to2125 h (Friedrichs et al., 1985; Hedden et al., 1986). The
aim was a conversion process from coal to SNG with significantly lower costs than the fixed-bed methanation concepts. As in the Bigas project, very high conversion wasreached within one reactor which allowed for significantly
lower equipment and capital costs (–30 %) than a series ofadiabatic fixed beds. Because of the falling oil and gas
prices, however, the development was stopped in the mid1980s.
Since about the year 2000, the development of this tech-

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nology has been pursued again, now for the implementation of wood gasification. Seemann et al. (2006) showedthat an internal regeneration of catalysts with carbon
depositions takes place in BFB methanation reactors operated with producer gas from wood gasification. Accordingly, the fluidized-bed reactor can be divided into different
zones. At the inlet zone, the deposition of carbon takes
place due to high concentrations of CO and unsaturatedhydrocarbons (ethene, benzene). In the zone above, steam
gasification and hydrogenation of the carbon deposits ortheir precursors are predominant. Ultimately, an equilibrium between carbon deposition and the regeneration of thecatalyst can be reached in this way.
Both for the conversion of wood gasification gas and forthe upgrading of biogas by methanation of the CO2 content
with added renewable hydrogen, extensive long-duration
field studies (>1000 h) with a container-based TRL 4/5
plant were conducted showing the catalyst’s resilienceagainst coke formation (Schildhauer and Biollaz, 2016;
Witte et al., 2019). As part of a long-term field test of over
1100 h regular operation, the particle elutriation rate ofcatalyst material was measured for a reactor with a diameter of 5.2 cm and an original catalyst mass of 800 g. This is
illustrated in Fig. 2, which shows the elutriation rates andthe concentration profiles over the duration of the field test.
Throughout the campaign, only 70 g of the catalyst was
elutriated from the reactor. The elutriation rate, measured
by an in-line PCME sensor, spiked when fresh catalystmaterial was inserted, because fine particles are elutriatedat much higher rates. It was therefore concluded that attrition rates are low and the catalyst material has to be replaced only once per year (Witte et al., 2019). Thesefindings are supported by the results of the experimental
campaigns for the Comflux process development, where
the fluidized bed methanation of coal gasification gas wasdemonstrated up to 20 MW in the early 1980s (Friedrichs
et al., 1985; Hedden et al., 1986). The concentration profiles were stable throughout the campaign and showed ahigh methane yield of 85–90 mol%. Further, optimizing thegas cleaning process step was identified in these tests to beof crucial importance for the extension of the catalyst lifetime, especially limiting the number of sulfur species(Calbry-Muzyka et al., 2019; Calbry-Muzyka andSchildhauer, 2020; Struis et al., 2009).
Other groups are now also investigating this technology,
for example in the French GAYA project (Hervy et al.,2021), in Korea (Nam et al., 2021) or in China; there, however, for the conversion of coal to methane (Liu et al.,2016).

4. Hydrodynamics of fluidized-bed reactors
with vertical internals

Understanding the hydrodynamics of BFB reactors is ofgreat importance for the design, scale-up and optimization

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100

Elutriation rate (g/h)

Elutriation rate [g/h]10Predictions u/umf = 2
Predictions u/umf = 5
Predictions u/umf = 8

1

0.1

0.01

0.001

0

200

400

600

800

1000

Molar fraction [%]Operational hours [h]Operational hours [h]Fig. 2 Long-duration test at TRL 5 (10 kW) with real biogas from the biogas plant in Zürich Werdhölzli. Results of the particle elutriation measurement and (dry) gas composition downstream of the reactor, adapted from Ref. (Witte et al., 2019).

of the reactor, especially in the case of the highly exothermic methanation process, where efficient heat removal isessential. Rising bubbles and correlated hydrodynamic
phenomena strongly influence key properties of catalytic
BFB reactors, e.g. the heat transfer between the bed andheat exchanger surfaces, gas–solid contacting, mass transfer between bubbles and the bed, residence time distributions, bed expansion, solids segregation and mixing as wellas catalyst lifetime, specifically the attrition and elutriationof catalyst particles (Glicksman, 1984; Grace, 2020;
Schildhauer and Biollaz, 2016; Valenzuela and Glicksman,
1985; Werther, 2007).
The most relevant bubble properties are the diameter db,
the rise velocity of the bubble ubr, the bubble hold-up εb andultimately the shape of the bubble, which allow a good estimation of the above-mentioned phenomena. Extensive
experimental studies have been conducted over several
decades; however, the majority of the correlations to predict bubble properties are empirical or semi-empirical andoften are only valid for a specific range of conditions.
Additionally, the increase of available computational
power in recent years led to significant advancement in thefield of computational fluid dynamics (CFD), which allows
the investigation of fluidized bed hydrodynamics throughnumerical simulations. While certain properties such as the
bubble hold-up εb, the average bed expansion or the minimum fluidization velocity umf can be predicted accurately
through CFD simulations, it usually requires empiricallydetermined tuning parameters to reach an agreement between simulated and experimentally measured bubble
properties. While many aspects of fluidized beds without
internals have been researched, there are far fewer publications on the influence of vertical heat exchanger tubes.
Simulations with CFD software (Liu and Hinrichsen, 2014;
Lv et al., 2020; Verma et al., 2016) were compared to experimental results and showed significant deviations between fluidized beds with and without vertical internals
where the internals such as vertical sheets or heat143

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exchanger tubes reduce the increase of the bubble size andlead to more elongated bubbles.

4.1 Pressure fluctuation measurements

Measuring the pressure drop and the pressure fluctuations caused by bubble passing, coalescence and bubble
eruptions is a simple, usually non-intrusive method that canbe applied in experimental studies on cold-flow models,
but also hot reactors (with some adaptations). It was foundthat vertical internals can decrease the pressure drop by up
to 10 % (Taofeeq and Al-Dahhan, 2018a), shift the onset ofslugging to higher gas flow rates (Coronella et al., 1994)
and also have an influence on the transition to turbulent
flow (Taofeeq and Al-Dahhanl, 2018b) in reactors with
10 cm and 14 cm inner diameter, respectively. For themeasurements of the pressure fluctuations with fastpiezo-electric probes, spectral decomposition was applied
to retrieve the average bubble size (Rüdisüli et al., 2012a;b; c). It could be shown that the bubbles in fluidized beds
with vertical tubes not only grow more slowly, but also
have a smaller hydraulic diameter. Thus, a better masstransfer between bubbles and the dense phase can be expected (Rüdisüli, 2012). This was recently confirmed for a
bed of densely packed, small-diameter tubes where thestandard deviation of the pressure fluctuations was halvedby the presence of the vertical tubes, while the diameter ofthe bubbles at the bed surface during eruptions went down
by a factor of around four (Antonini, 2019).
Horizontal heat exchanger tubes with the necessary
packing density, on the other hand, led to gas cushions andlocal defluidization, which means that relatively poor masstransfer can be expected (Rüdisüli et al., 2012d). In addition, the transversely installed pipes should limit the vertical transport of the particles, which counteracts thelimitation of carbon deposits by internal regeneration.

4.2 Sector approach

Rüdisüli et al. (2012d) also presented a modified approach for the up-scaling of fluidized-bed reactors withvertical internals, which considers defined sectors of internals as autonomous regions. As was already assumed by
Volk et al. (1962), it was shown that the vertical pipes, witha spacing of only a few centimeters, strongly structure thefluidized bed and thus minimize the influence of the reactor

wall. Experiments with optical probes were conducted withvariation of the cross-section of a fluidized bed cold-flow
model while keeping tube diameter and pitch constant (i.e.
by increasing the number of tubes in the same arrangement). From at least two complete rows of tubes around ameasuring point, the measurement is largely representative
for all other such “elementary cells” and can be used todesign larger reactors with a correspondingly increased
number of vertical tubes (with the square of the diameter
ratio) (Maurer et al., 2014), see also Fig. 3. This significantly simplifies the construction of pilot plants and theresulting scale-up, as was confirmed by comparing thepressure fluctuations from a pilot scale (1 MW) plant inGüssing and a bench-scale cold-flow model (Rüdisüli et al.,2012d).

4.3 Optical probes incl. Monte Carlo

Optical probes are versatile tools to facilitate deeper insight into the hydrodynamics of fluidized beds. While thepressure fluctuation measurements only give informationon the average size of the rising bubbles, a vertical arrangement of two optical probes can deliver the rise velocity of abubble, which in combination with the signal duration canbe translated into a pierced chord length. The challenge is,however, that the optical probes only allow local measurements, and thus many radially and axially resolved measurements as well as several statistical corrections areneeded. Statistical analysis and a Monte Carlo study led
Rüdisüli et al. (2012e) to the conclusion that the mean
chord length measured with optical probes in freely bubbling beds is a good approximation for the mean bubble
diameter. This is due to the increased pierce probability oflarger bubbles (which lead to an increased average chord
length), which is counterbalanced by the fact that the optical probes do not necessarily pierce the bubbles at theirlongest axial expansion. This was confirmed by Maurer etal. (2015a) who, however, found that these two counteracting effects are not valid for the measurement of the mean
bubble rise velocity for which a different correction approach is required. Further measurements with opticalprobes in fluidized beds with vertical internals showedsignificantly different chord lengths and rise velocities(Taofeeq and Al-Dahhan, 2018c). It is therefore necessary
to adjust the data evaluation appropriately.

Tradi�onal
scale-upØ 14 cm

Fig. 3144

dInternals,small
dInternals,large

Review Paper

Sectoralscale-up= 1:3
Ø 22 cmdInternals,small
dInternals,large= 1:1

Sectorial scaling approach for fluidized bed reactors with vertical internals.

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Light path from laser source
into the reactor
Light path from reactor
to photo sensors
Details of optical probe.

u/umf [-]

ubr [m/s]

Fig. 4

Chord length [cm]Fig. 5Results (right) of optical probe measurements in a pressurized pilot scale plant (left) at a pressure of 6 bar and a measurement height of 32 cm.

While most cold-flow models reported in the literature
are operated at ambient pressure, Maurer (2015) developed
and tested a new optical probe sensor, which is able tomeasure under reactive conditions with pressures of up to
2.5 × 106 Pa and temperatures of up to 400 °C, see Fig. 4.
First cold-flow measurements in the pilot plant GanyMeth (22.4 cm inner diameter, operation up to
1–11 × 105 Pa, see Fig. 5, left) with pressurized air andalumina particles at pressures of up to 5 × 105 Pa wereconducted and proved the system (Schillinger, 2018).
Recently, a unique data set was measured with these radially movable optical probes at six heights and six radial
positions for Geldart-A and Geldart-B-type bed material,pressures of 1–11 × 105 Pa and fluidization numbers u/umf
of 2–10, see Fig. 5 (right) (Riechmann, 2022). It becomes
obvious that not a single value but rather distributions ofbubble properties (chord length, rise velocity ubr) exist foreach experimental setting. Meanwhile, data evaluation is
ongoing to obtain correlations for the distributions of bubble properties to enable improved modelling/simulation ofsuch reactors.

4.4 X-ray tomography

In cooperation with TU Delft, X-ray tomographic investigations were conducted and evaluated by an algorithm
allowing the reconstruction of the bubbles in a fluidized
bed of Geldart-B particles. This allowed determining theirproperties with regard to spatial distribution over the
cross-section, bubble size and bubble velocity (Maurer etal., 2015b). As can be seen in Fig. 6, a square or concentricarrangement of heat exchanger tubes has very similar effects: there are more bubbles compared to the configuration
without internals, and the bubbles deviate significantly
from the round shape. Further, experiments with cold-flow
models of two diameters (14 cm and 22 cm) confirmed
once again the sectorial scaling approach: the differences
between beds of different sizes are small if they are structured by vertical internals, while in fluidized beds without
internals, the bubble size increases significantly with thereactor diameter.
Compared to measurements with local samples (e.g. optical sensors), the possibility of X-ray tomography to determine the bubble size and the bubble velocity of all risingbubbles over the entire cross-section is a great advantage,

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Fig. 6 Reconstructed passing of bubbles in the measurement height (chronological order, z-axis = time) for beds without (left) and with internals(middle, right) (Schillinger, 2018).

as such measurements are less subject to statistical effects.
To derive conclusive bubble property distributions from thelocal measurement at a measurement height, measurements
must be taken at several radial positions, assuming rotational symmetry. Bubbles with a larger cross-section in the
direction of flow have a higher probability of hitting thesensor, but the bubbles are not always hit at the point of thelongest expansion of the bubble in the direction of flow as
already mentioned above.
X-ray tomography reconstructions were also used tosimulate the measurements of an ideal virtual opticalprobe. The resulting chord lengths and rise velocities werethen compared to the bubble properties derived from the
X-ray tomography reconstructions. Statistical effects resulting from the assumed shape of the bubbles, such as
chord length distributions and the reduced probability ofthe optical probe to pierce elongated bubbles in comparison
to oblate bubbles, were identified and correction factors
were proposed accordingly (Schillinger, 2018). This approach opens the door to improving the evaluation of local
optical probe data in ongoing work that is additionally
supported by projecting X-ray radiography. Neutron radiography (Ozawa et al., 2002) as well as recent X-ray radiography data (see Fig. 7) show not only bubble sizes andtheir rise velocities but also the interactions with opticalprobes. Further, the trajectories can be observed whichshow that bubbles rise vertically in between or if larger,
along the vertical tubes, while in freely bubbling beds significantly more lateral movements and therefore increased
coalescence is found (Ozawa et al., 2002).

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4.5 Heat transfer

The importance of removing or supplying heat is themain motivation for introducing vertical heat exchanger
tubes into fluidized beds, as they allow reaching more orless isothermal operation conditions. For the proper designof the reactors, quantitative data of heat transfer are needed.
Already in 1996, Gunn and Hilal (1996) measured the totalheat transfer from electrically heated tubes to the bed ofglass beads (BALLOTINI) at ambient pressure. For
Geldart-A-type material, total heat transfer coefficientsfrom 350 W/m2K (at u/umf = 2) to 550 W/m2K (at u/
umf > 7) were found, independent of the number of tubes
and the gas distributor. For Geldart-B materials, the best
results were reached with a porous-plate-type distributor,
ranging from 250 W/m2K (at u/umf = 2) to >350 W/m2K (atu/umf > 5). Ozawa et al. (2002) measured total heat transfer
coefficients from 250 W/m2K (at u/umf = 2) to 400 W/m2K(at u/umf > 5) for Geldart-B-type alumina particles at ambient pressure.
These results are confirmed by our own measurements of
PSI at the pilot scale plant GanyMeth, heating up cold nitrogen at 4 × 105 Pa with thermo-oil of 250 °C in a bed of
Geldart-B alumina particles. At a fluidization number ofaround 1.3, the axial temperature profile from the distributor to 1.5-m height ranged from 232 °C to 243 °C, proving
the strong axial heat dispersion in these systems. The estimated total heat transfer coefficient (though not very precise due to the small temperature difference of 7 to 18 K) is
170 W/m2K, which fits exactly to the results of (Gunn and
Hilal, 1996) for Geldart-B materials and a porous plate

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u/umf = 3

u/umf = 4

Review Paperu/umf = 6

Without
internals

Withinternals

Fig. 7 Recent X-ray radiography data (PSI in collaboration with TU Delft) in a 22.4 cm column at ambient pressure with Geldart-B-type material(γ-alumina particles).

H = 40 cm; Twater,in = 40°C; TAir = 20°C
300.0

0.20
α R0

0.18

α R3
Bubble Holdup R0

α [W/m²K]

200.0

0.14

Bubble Holdup R3

0.12
0.10

150.0

0.08

100.0

0.06
0.04

50.0
0.0

Fig. 8

R3

0.16

R0

Bubble Holdup εb [-]

250.0

0.02
1.5

2

u/umf

3

4

0.00

Results of locally applied optical and heat transfer probes in a cold-flow model (diameter 22.4 cm) with Geldart-B-type alumina particles.

distributor at such low fluidization.
Taofeeq et al. introduced advanced non-invasive fast response heat transfer probes in the research field and usedthem for local measurements in a 14 cm column with
Geldart-B-type glass particles for fluidization numbers
from 1.4 to 2.5 (Taofeeq and Al-Dahhan, 2018a, b, c;
Taofeeq et al., 2022). For fluidization numbers below 2,
they found that the presence of vertical internals improved
the local heat transfer by up to 20 % in regions closer to thewall, but not in the center of the column. Generally, local
heat transfer coefficients from 270 W/m2K (at u/umf = 1.4)
to 320 W/m2K (at u/umf = 2.5) were found for beds with an
aspect ratio of >1.5.

Further, they found that the local heat transfer is enhanced by the presence of bubbles; our own measurements
at PSI with a similar approach with locally applied opticaland heat transfer probes (Geldart-B-type alumina particle
in a cold-flow model) showed even a direct proportionalityof bubble hold-up and local heat transfer, see Fig. 8.

5. Conclusions and outlook

Bubbling fluidized bed (BFB) reactors with immersed
vertical heat exchanger tubes proved to be a suitable reactor type for strongly exothermic reactions such as themethanation of CO/CO2-rich gases to form synthetic natural gas. They help to avoid hotspots and the risk of thermal

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reactor runaway, potentially leading to catalyst deactivation.
Many studies in literature and our own experiments wereable to show that the catalyst particle movement does not
only allow for optimal heat removal, close to isothermal
operation and thus little catalyst sintering, but that the catalyst particle movement over the height of the reactor withdifferent concentration zones favors the chemical
catalyst stability. This contribution reviewed the fluiddynamic studies for BFB reactors with immersed heatexchangers in the last decades comprising studies withpressure fluctuation probes, optical probes, X-ray tomography and studies on heat transfer.
With the results obtained so far, it can be concluded thatimmersed vertical heat exchanger tube banks improve thefluidization and have the potential to facilitate up-scaling.
Further, important knowledge was collected to better describe and correlate the distribution of sizes and rise velocities of bubbles which improve heat transfer and particle
movement but may limit the reactor performance in thecase of too low mass transfer. With the obtained correlations, the reactor models can be significantly improved interms of mass transport predictions. Planned measurementsof temperature and concentration profiles during reactive
operation in PSI’s pilot plant will help to practically validate the methanation reactor model in the experimentallyinvestigated pressure and temperature range. This together
with the sector scaling model will allow upscaling to thedifferent applications with significantly reduced risk.

Acknowledgments

The authors would like to thank Marion Ziegler, Roham
Seif, Evert Wagner, Ruud van Ommen and the whole TCPteam at PSI for support during experiments and for scientific discussions. Further, the financial support of the Energy Systems Integration platform at PSI is acknowledged.

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Authors’ Short Biographies
Philipp Riechmann
Philipp Riechmann studied process and energy engineering at TU Berlin and is a scientist at thePaul Scherrer Institut with focus on the hydrodynamic aspects of fluidized beds. He recently defended his PhD thesis successfully.

Tilman Schildhauer
Tilman Schildhauer is Senior Scientist at Paul Scherrer Institut in Switzerland where he leads theresearch in methanation, gas cleaning and PtX processes. He studied chemical engineering in Dortmund, received a PhD from ETH Zurich and spent his postdoc time in the Reactor and Catalysis
Engineering group at TU Delft.

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