Lab scale biogas reformer D5.4 First model of novel ... · Start date of project: 01/09/2015...

BIONICO

BIOGAS MEMBRANE REFORMER FOR DECENTRALIZED HYDROGEN PRODUCTION

FCH JU GRANT AGREEMENT NUMBER: 671459

Start date of project: 01/09/2015 Duration: 3 years

WP5 – Lab scale biogas reformer

D5.4 First model of novel reforming membrane reactor system Topic: FCH-02.2-2014 - Decentralized hydrogen production from clean CO2-containing biogas Type of Action: FCH2-RIA Research and Innovation action Call identifier: H2020-JTI-FCH-2014-1

Due date of deliverable: 2016-08-31

Actual submission date: 2016-09-07

Reference period: 2015-09-01-2016-09-07

Document name: BIONICO-WP5-D51-DLR-TuE-20160907-v02.doc Prepared by (*): TUE

Version DATE Changes CHECKED APPROVED

V1 2016-08-23 First release TUE Niek de Nooijer

V2 2016-09-07 Final version TUE FGallucci

Dissemination Level

PU Public X

PP Restricted to other programme participants (including the Commission Services)

RE Restricted to a group specified by the consortium (including the Commission Services)

CO Confidential, only for members of the consortium (including the Commission Services)

CON Confidential, only for members of the Consortium

This project has received funding from the Fuel Cells and Hydrogen 2 Joint Undertaking under grant agreement No 671459. This Joint Undertaking receives support from the European Union’s Horizon 2020 research and innovation programme and Hydrogen Europe and Hydrogen Europe Research.

___________________________________________________________________________________ (*) indicate the acronym of the partner that prepared the document

D5.4 First model of novel reforming

membrane reactor system

Proj. Ref.: BIONICO - 671459 Doc. Ref.: BIONICO-WP5-D55-DLR-TUE-20160907-v02 Date: 2016/09/07 Page Nº: 2 of 13

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PUBLISHABLE SUMMARY A first modelling is performed for the fluidized bed membrane reactor for biogas steam reforming. A one dimensional two-phase fluidization model is considered for the simulation of fluidized bed membrane reactor. Some first base case simulations results show that the system is mass transfer limited, therefore to accurately simulate the lab scale reactor the mass transfer is of high importance to be well described. Furthermore the influence of the presence of the dry reforming reaction is studied. Currently the kinetics for the catalyst used in the final process is still being studied. A rate equation for DR is assumed to describe the dry reforming kinetics. The results show that the production of hydrogen in the first part of the reactor increases however the overall equilibrium conversion of methane is lowered. For the final model the experimentally obtained kinetics will be implemented in the model. Furthermore a description of the mass transfer resistance near the membrane will be included. To validate the model experimental results of a single membrane reactor and the lab scale reactor will be used.




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Content

PUBLISHABLE SUMMARY .................................................................................................................. 2

1. EXECUTIVE SUMMARY ............................................................................................................... 4

1.1. Description of the deliverable content and purpose ................................................................. 4

1.2. Brief description of the state of the art and the innovation brought ........................................ 4

1.3. Deviation from objectives ............................................................................................................ 4

1.4. If relevant: corrective actions ...................................................................................................... 4

1.5. If relevant: Intellectual property rights ....................................................................................... 4

2. INTRODUCTION ............................................................................................................................ 5

3. DESCRIPTION OF THE MODELS................................................................................................. 7

3.1. Kinetic model ............................................................................................................................... 7

3.2. Membrane model .......................................................................................................................... 8

3.3. Hydrodynamics and mass transfer model ................................................................................. 8

4. RESULTS ...................................................................................................................................... 10

5. CONCLUSIONS ............................................................................................................................. 12

REFERENCES ...................................................................................................................................... 12




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1. EXECUTIVE SUMMARY 1.1. Description of the deliverable content and purpose The main task reported in this document is a first model approach and model results of the fluidized bed membrane reactor, for steam methane reforming and autothermal reforming of biogas, for the production of hydrogen. The model discussed in this report is a phenomenological model of the system, which incorporates information of the developed membranes fluxes and selectivity and the catalysts micro kinetics. The final model will be used to scale up and design the novel reactor system. The final model will be validated with the lab scale experimental results and results from the pilot plant. 1.2. Brief description of the state of the art and the innovation brought First phenomenological model analysis of the membrane assisted fluidized bed reactor for steam reforming and autothermal reforming of biogas. 1.3. Deviation from objectives N/A 1.4. If relevant: corrective actions N/A 1.5. If relevant: Intellectual property rights TUE on modelling.




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2. INTRODUCTION In this report, a first overall assessment of model of the fluidized bed membrane reactor (FBMR) for

steam reforming and auto-thermal reforming of biogas and combined H2 separation is discussed. The

schematic of the membrane reactor is represented in Figure 1.

Figure 1: Membrane Assisted Fluidized Bed Reactor configuration

The FBMR is a system that combines the reforming reactions and the separation of hydrogen in one

single reaction unit. Palladium based membranes are immersed in the fluidized bed of catalytic particles.

The membranes are able to separate the produced hydrogen. In this way the equilibrium of the reactions

are shifted and higher conversions are obtained, furthermore a pure stream of hydrogen is obtained. The

system is fed with biogas and steam or steam and air. Air is mixed with H2O in order to reach the

required steam-to-carbon SCR and oxygen-to-carbon OCR which are calculated as follow.

𝑆𝐶𝑅 =𝐹𝐻2𝑂

∑ 𝑖 ∙ 𝐹𝐶𝑖𝐻2𝑖+24𝑖=1

(1)

𝑂𝐶𝑅 =𝐹𝑂2

∑ 𝑖 ∙ 𝐹𝐶𝑖𝐻2𝑖+24𝑖=1

(2)

The air+H2O are pre-heated up to the inlet temperature (450-550°C depending on the operating

conditions of the FBMR) and fed to the system. In the first part of the reactor, the CH4 reacts with O2 and

H2O to produce H2O, CO and CO2. Most of the O2 is consumed and some H2 is produced due to the

reforming and CH4 partial oxidation reaction. The reactor is operated at 3-12 bar. the produced H2

permeates at to the permeate side where the H2 partial pressure is lower due to the presence of sweep

gas or vacuum conditions.

In this report a sensitivity analysis of the membrane assisted autothermal reforming (MA-ATR)

performance is carried out and different designs are compared. An important parameter described is

H2

CO2+H2O+N2+traces of (CH4+H2+CO)

catalyst

Pd-Ag membrane layer

CH4+H2O+air CH4+H2O+air

MA-FBR/vacuum




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Hydrogen Recovery Factor (HRF) which is defined as the ratio between permeated hydrogen and the

maximum flow of hydrogen that could be produced. It is calculated as follows:

𝐻𝑅𝐹 =𝐹𝐻2,𝑝𝑒𝑟𝑚

𝐹𝐻2,𝑒𝑞𝑢𝑖𝑣𝑎𝑙𝑒𝑛𝑡

𝐹𝐻2,𝑒𝑞𝑢𝑖𝑣𝑎𝑙𝑒𝑛𝑡1 = ( ∑ 𝐹𝑖

𝐻𝑛𝐶𝑚

𝑖

× (2𝑛𝑖 +𝑚𝑖

2) + 𝐹𝐶𝑂 + 𝐹𝐻2

− 2 × 𝐹𝑂2)

𝑓𝑒𝑒𝑑

The first model will represent the lab scale reactor where two scales are considered, single membrane and 5 membranes. The main assumptions and parameters behind the first model are summarized in the table below:

Parameter units value

Feed & operating conditions

Model Biogas: different compositions representing biogas from landfill and from an anaerobic digester.

Uniform reforming temperature °C 600 (500-600)

Pressure reaction side bar 8 (3-12)

Pressure permeate side bar 0.3

SCR - 2.0-5.0

OCR - 0.2-0.4

Membranes - 1-5

CO2/CH4 - 0.3-0.8




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3. DESCRIPTION OF THE MODELS The analysis of the system is based assuming the system under steady-state conditions and therefore

any dynamic operations have been considered. The analysis is based on a pseudo-homogeneous one-

dimensional model in which: i) any radial effect in the temperature profile; ii) the gas are perfectly mixed

and there is no diffusion occurring from the bulk phase to the membrane surface (concentration

polarization are neglected); iii) no diffusion limitations are occurring inside the catalyst.

3.1. Kinetic model The kinetic model used for the analysis should take the steam methane reforming (SMR) and water gas

shift (WGS) into account. Due to the presence of CO2 in the feed gas, the dry reforming (DR) can play a

significant role. In the case of ATR methane combustion (MC) is also taken into account.

SMR: 𝐶𝐻4 + 𝐻2𝑂 ⇄ 3𝐻2 + 𝐶𝑂

WGS: 𝐶𝑂 + 𝐻2𝑂 ⇄ 𝐻2 + 𝐶𝑂2

DR: 𝐶𝐻4 + 𝐶𝑂2 ⇄ 2𝐻2 + 2𝐶𝑂

MC: 𝐶𝐻4 + 2𝑂2 → 2𝐻2𝑂 + 𝐶𝑂2

The kinetics used in this model will be based on the kinetics found by Marra et al. for a RhZrO2 catalyst

[1]. This model uses the Kinetic expressions for Numaguchi and Kikuchi (NK) [2] to describe SMR and

WGS, Trimm and Lam [3] is used to describe the methane combustion. The DR reforming reaction is not

taken into account in the first model. The influence of the dry reforming reaction on the kinetics is being

studied and will be included in the final model.

The reaction rates R [mol m-3s-1] of the abovementioned reactions are calculated as follows (

Table 1):

Table 1: kinetic model expressions

Reaction rates expressions

. 4 2 . 4 2

2

4 4 2 24 4 2 211

a MC CH O b MC CH OMC ox oxox ox

CH CH O OCH CH O O

k p p k p pR

K p K pK p K p

3

4 2 2 ,

1.596

2

/SMR CH H O H CO eq SMR

SMR

H O

k p p p p KR

p

2 2 2 ,

2

/WGS CO H O H CO eq WGS

WGS

H O

k p p p p KR

p

Pre exponential factor

(mol/kg s bar2) Activation energy (kJ/mol)

1 1 2

, ( bar )a MC catk mol s kg 56.7 ·105 86.0 1 1 2

, ( bar )b MC catk mol s kg 47.7 ·105 86.0 1 1 0.404( bar )SMR catk mol s kg 9.7·104 83.6

1 1 1( bar )WGS catk mol s kg 17.2 ·102 54.5 1

4 ( )ox

CHK bar 1.26 ·10-1 -27.3 1

2 ( )ox

OK bar 7.87 ·10-7 -92.8




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,

,

20009ln 22.82

( )

4400ln 4.036

( )

eq SMR

eq SMR

KT K

KT K

3.2. Membrane model

The expression used for hydrogen permeation is described in equation (4) where “𝑃𝐻2” is the pure

hydrogen permeance value (mol/s.m2.Pan), the term inside brackets represents the driving force for the

permeation, and “𝑛” is the exponential factor. In this work, the parameters of an asymmetric membrane,

developed by Tecnalia, with a Pd based layer of 4.5 μm on a ceramic support, are implemented. [4]. The

initial production of hydrogen takes place in the inlet section of the reactor prior to the membranes (also

referred to as pre-reforming), then in the membrane reactor section, the extraction of hydrogen through

the perm-selective membrane drives the equilibrium reaction towards the products.

𝐹𝐻2,𝑝𝑒𝑟𝑚 = ∫ 𝑃𝐻2∙ (𝑝𝐻2,𝑟𝑒𝑡

𝑛 − 𝑝𝐻2,𝑝𝑒𝑟𝑚𝑛)𝑑𝐴

𝐴𝑚𝑒𝑚

(3)

The The H2 flux (𝜙𝐻2" ) has been calculated using the following equation, according to experimental data:

𝜙𝐻2" (𝑚𝑜𝑙 𝑚−2𝑠−1) =

𝑃𝑚

𝑡(𝑝𝐻2,𝑟𝑒𝑡

𝑛 − 𝑝𝐻2,𝑝𝑒𝑟𝑚𝑛 )

𝑃𝑚

𝑡= 𝑃0𝑒−

𝐴𝑒𝑅𝑇

Pm is described as function of temperature (6) with P0 and the activation energy of the membrane given in Table 2 together with the further membrane details. The membrane parameters are obtained from experimental data.

Table 2: main membrane parameters

Membrane n Tickness P0 (mol/m2/sec/Pa^n) Ea (kJ)

PdAgAu 6% (Al2O3 100nm 14/7 OD/ID)

0.525 - (included in P0) 8.58 x 10-3 8.92

3.3. Hydrodynamics and mass transfer model The membrane assisted fluidized bed model merged the concept of reaction kinetics with bed

hydrodynamics and gas separation. A one dimensional two-phase fluidization model is considered for

the simulation of fluidized bed membrane reactor. This model has been developed in the past years and

used for different process involving H2/O2 membrane reactors [5,6]. There are two phases considered in

the reactor i.e. bubble and emulsion phase. The schematic of the reactor is a combination of different

continuous stirred tank reactors (CSTRs) each section consists of further sub CSTRs. The bubble phase

is simulated as a pseudo-PFR and therefore the number of CSTR is increased as shown in Figure 2.

The H2 permeates through the membrane from both the emulsion and the bubble phase, which however

are working with a different gas composition due to the mass transfer of the system. For the MA-ATR,

the energy balance is accounted only for the overall system and it is not implemented in the model. For

this reason, the analysis is carried out at iso-thermal condition (550 °C).




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Figure 2: Schematic of the MA-FBR model

Moreover, the model also assumes that the gas passing through the emulsion phase is at minimum

fluidization velocity and is properly mixed in each section while the remaining gas flows as bubbles. The

gas permeated from both bubble and emulsion phase through membrane is distributed on the basis of

local bubble fractions. Meanwhile, the gas separated from emulsion phase is immediately compensated

through bubble phase (to keep the emulsion phase at minimum fluidization). The gas compensated

through bubbles phase is depends on the value of bed voidage (bed expansion). The bed voidage (ε)

represents the division of membrane area between bubble and emulsion phase. The bubbles are free of

catalyst, both the reactions i.e. SMR and WGS are happening in the emulsion phase. A constant

temperature is considered across the reactor bed with no heat loses to the surroundings. The

correlations and equations used in the modelling are shown in Table 3.

Table 3: Hydrodynamic parameters used in the modelling

Parameters Equation Ref.

Archimedes number 𝐴𝑟 = 𝑑𝑝3𝜌𝑔(𝜌𝑝 − 𝜌𝑔)𝑔/𝜇𝑔

2 [7]

Minimum fluidization velocity 𝑢𝑚𝑓 = (𝜇𝑔 𝑑𝑝𝜌𝑔⁄ ) (√(27.2)2 + 0.0408𝐴𝑟 − 27.2) [8]

Bed voidage at minimum fluidization velocity 휀𝑚𝑓 = 0.586𝐴𝑟−0.029 (𝜌𝑔

𝜌𝑝)

0.021

[8]

Velocity of rise of swarm of bubbles 𝑢𝑏 = 𝑢𝑜 − 𝑢𝑚𝑓 + 𝑢𝑏𝑟

[7]

Rising velocity of single bubble 𝑢𝑏𝑟 = 0.711(𝑔𝑑𝑏,𝑎𝑣𝑔)1/2 [7]

Emulsion velocity 𝑢𝑒 =𝑢𝑜 − 𝛿𝑢𝑏

1 − 𝛿 [7]

Average bubble diameter 𝑑𝑏,𝑎𝑣𝑔 = 𝑑𝑏,𝑚𝑎𝑥 − (𝑑𝑏,𝑚𝑎𝑥 − 𝑑𝑏𝑜)exp ( −0.3𝐻

𝐷𝑇) [9]

Initial bubble diameter 𝑑𝑏𝑜 = 0.376(𝑢𝑜 − 𝑢𝑚𝑓)2

[6]

Bubble phase fraction 𝛿𝑏𝑛 =𝑢𝑏

𝑠

𝑢𝑏 [6]

Emulsion phase fraction 𝛿𝑒𝑛 = 1 − 𝛿𝑏𝑛 [6]

Maximum superficial bubble gas velocity 𝑢𝑠

𝑏,𝑚𝑎𝑥 = 𝑢𝑜 − 𝑢𝑚𝑓 [6]

Initial superficial bubble gas velocity 𝑢𝑠

𝑏,𝑜 = 𝑢𝑏𝑟,𝑜𝛿𝑏𝑜

where 𝛿𝑏𝑜 = (1 − 𝐻𝑚𝑓 𝐻𝑓⁄ ) [6]

Height of bed at minimum fluidization velocity 𝐻𝑚𝑓 = 𝐻𝑠

1 − 휀𝑠

1 − 휀𝑚𝑓

[9]

EE

B B B B

Permeated H2,e

Perm. H2,b

Bubble-to-emulsion mass and energy

exchange

Bubble-to-emulsion mass and energy

exchange

Feed gas

Perm. H2,b

Permeated H2,e

Perm. H2,b Perm. H2,b

retentate




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Height of bed expansion

𝐻𝑓 = 𝐻𝑚𝑓

𝐶1

𝐶1 − 𝐶2

[6]

where,

𝐶1 = 1 −𝑢𝑏,𝑜

𝑢𝑏,𝑎𝑣𝑔𝑒𝑥𝑝 (−

0.275

𝐷𝑇)

𝐶2 =𝑢𝑏

𝑠

𝑢𝑏,𝑎𝑣𝑔[1 − 𝑒𝑥𝑝 (−

0.275

𝐷𝑇)]

Average bubble rise velocity 𝑢𝑏,𝑎𝑣𝑔 = 𝑢𝑜 − 𝑢𝑚𝑓 + 0.711(𝑔𝑑𝑏,𝑎𝑣𝑔)1/2 [6]

Gas exchange coefficient

𝐾𝑏𝑐 = 4.5 (𝑢𝑚𝑓

𝑑𝑝) + 5.85 (

𝐷𝑔1/2𝑔1/4

𝑑𝑏5/4 )

[6] 𝐾𝑐𝑒 = 6.77 (

𝐷𝑔휀𝑚𝑓𝑢𝑏

𝑑𝑏3 )

1/2

1

𝐾𝑏𝑒=

1

𝐾𝑏𝑐+

1

𝐾𝑐𝑒

4. RESULTS The analysis has been carried by changing the total inlet flow changing u/umf. In terms of H2 extraction

and methane conversion where the choice of different u/umf. The following simulations have been

performed introducing first the kinetic rate limitations (KL), then the emulsion-to-bubble mass transfer

limitations (MTL). The reactor design implemented in the 1D two-phase model consider the presence of

5 membranes with an outer diameter of 0.014 m and a length of about 0.2 m. The figure points out that

the kinetic rate limitations do not have a big impact on the reactor performances since the oxidation and

reforming reactions are however faster in compared to the mass transfer limitation of each species.

Results of the phenomenological model show an H2 recovery flow of 96% at u/umf 2.5. The maximum H2

extraction value cannot be taken into account because the inlet feed velocity is not high enough to

sustain the fluidization regime along the entire membrane reactor. The decrease in hydrogen recovery

after 4 times Umf is due to the membrane flux that becomes limiting.

Figure 3 Left figure shows the conversion as function of U0/Umf, in the right figure the Hydrogen recovery factor is given as a fuction of U0/Umf.The case of natural gas (NG) is compared with biogas (BG)

In Figure 3 also the case of biogas is included (right). For the biogas case in Figure 4 the gas

composition along the membrane reactor founded by the phenomenological model at u/umf equal to 3.

The hydrogen fraction increases at the beginning due CH4 conversion and after that it decreases along




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the reactor due to the hydrogen permeating through the membrane. Therefore the equilibrium is shifted

and more methane is converted.

Figure 4 example of reactant composition along the membrane reactor, at 500°C, 3 U0/Umf, SCR 3 and CO2/CH4 0.5.

As mentioned before, this model only takes into account only the kinetics of the SMR and WGS. To see the influence of the dry reforming a theoretical rate for dry reforming is implemented. The results of the composition along the reactor are shown in Figure 5. The dry reforming kinetics are assumed to be the same as for SMR however with the equilibrium constant for DR resulting in the following equation:

2 2 2 ,

1.596

2

/SMR CO H O H CO eq DR

SMR

H O

k p p p p KR

p

0

0.05

0.1

0.15

0.2

0.25

0.3

0.35

0.4

0.45

0.5

0 0.02 0.04 0.06 0.08 0.1 0.12 0.14 0.16 0.18 0.2

Frac

tio

n [

-]

Lenght [m]

CH4 H2O H2 CO CO2 N2




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Figure 5 example of reactant composition along the membrane reactor, at 500°C, 3 U0/Umf, SCR 3 and CO2/CH4 0.5, including theoretical dry reforming kinetics

From the results can be seen that the production of hydrogen is initially increased together with the conversion of methane as a result of a combination DR and SMR. The kinetics implemented in this first model are however a first estimate. To improve the model the experimentally obtained kinetics are required to be implemented.

5. CONCLUSIONS A first modelling is performed for the fluidized bed membrane reactor for biogas steam reforming. A one dimensional two-phase fluidization model is considered for the simulation of fluidized bed membrane reactor. First base case simulations results show that the system is mass transfer limited, therefore to accurately simulate the lab scale reactor the mass transfer is of high importance to be well described. Furthermore the influence of the presence of the dry reforming reaction is studied. Currently the kinetics for the catalyst used in the final process is still being studied. A rate equation for DR is assumed to describe the dry reforming kinetics. The results show that the production of hydrogen in the first part of the reactor increases. For the final model the experimentally obtained kinetics will be implemented in the model. Furthermore a description of the mass transfer resistance near the membrane should be included. To validate the model experimental results of a single membrane reactor and the lab scale reactor will be used. REFERENCES

[1] L. Marra, P.F. Wolbers, F. Gallucci, M.V.S. Annaland, Development of a RhZrO 2 catalyst for low temperature autothermal reforming of methane in membrane reactors, Catal. Today. 236 (2014) 23–33. doi:10.1016/j.cattod.2013.10.069.

[2] K. Numaguchi, T, Kikuchi, Intrinsic kinetics and design simulation in a complex reaction network; steam-methane reforming, Chem. Enginnering Sci. 43 (1988) 2295 – 2301.

[3] D.L. Trimm, C.W. Lam, The combustion of methane on platinum-alumina fibre catalysts—I. Kinetics and mechanism, Chem. Eng. Sci. 35 (1980) 1405–1413.

0

0.05

0.1

0.15

0.2

0.25

0.3

0.35

0.4

0.45

0.5

0 0.02 0.04 0.06 0.08 0.1 0.12 0.14 0.16 0.18 0.2

Frac

tio

n [

-]

Lenght [m]

CH4 H2O H2 CO CO2 N2




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[4] E. Fernandez, A. Helmi, K. Coenen, J. Melendez, J.L. Viviente, D.A.P. Tanaka, et al., Development of thin Pd–Ag supported membranes for fluidized bed membrane reactors including {WGS} related gases, Int. J. Hydrogen Energy. 40 (2015) 3506–3519. doi:http://dx.doi.org/10.1016/j.ijhydene.2014.08.074.

[5] F. Gallucci, M. Annaland, J. Kuipers, Autothermal reforming of methane with integrated CO2 capture in a novel fluidized bed membrane reactor. Part 1: experimental demonstration, Top. Catal. 51 (2008) 133–145. doi:10.1007/s11244-008-9126-8.

[6] F. Gallucci, M. Van Sint Annaland, J.A.M. Kuipers, Autothermal Reforming of Methane with Integrated CO2 Capture in a Novel Fluidized Bed Membrane Reactor. Part 2 Comparison of Reactor Configurations, Top. Catal. 51 (2008) 146–157. doi:10.1007/s11244-008-9127-7.

[7] D. Kunii, O. Levenspiel, Fluidization Engineering, Elsevier, 1991. doi:10.1016/B978-0-08-050664-7.50012-3.

[8] C.-Y. Shiau, C.-J. Lin, Equation for the superficial bubble-phase gas velocity in fluidized beds, AIChE J. 37 (1991) 953–954. doi:10.1002/aic.690370619.

[9] S. Mori, C.Y. Wen, Estimation of bubble diameter in gaseous fluidized beds, AIChE J. 21 (1975) 109–115. doi:10.1002/aic.690210114.

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