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Eindhoven University of Technology
MASTER
Autothermal reforming of methane in membrane assisted reactors
Ottenheijm, I.N.
Award date:2015
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Process Engineering
Multiphase Reactors group (SMR)
Department of Chemical Engineering and
Chemistry
Den Dolech 2, 5612 AZ Eindhoven
P.O. Box 513, 5600 MB Eindhoven
The Netherlands
www.tue.nl
Author:
I.N. Ottenheijm
ID: 0657126
Graduation Committee:
prof.dr.ir. M. van Sint Annaland
dr. F. Gallucci
dr.ir. E. Zondervan (external)
Ir. K. Coenen
MSc thesis
Date
28 October 2014
Autothermal reforming of methane
in membrane assisted reactors
I.N. Ottenheijm
October 2014 - Confidential
Technische Universiteit Eindhoven University of Technology
I Abstract
Abstract
Hydrogen is a versatile energy carrier which can be used in numerous industries. When used for
energy, the only product is water vapor, making it very friendly for the end-user. Hydrogen is
produced mainly from fossil fuels of which 48% is natural gas. Steam reforming of methane is a
process to produce hydrogen from methane and is currently performed in a multistage process at
elevated temperatures up to 900 °C to favor the thermodynamic equilibrium towards hydrogen.
In this work, steam methane reforming is combined with methane oxidation to achieve autothermal
operation: autothermal reforming of methane. The multistage process is combined in a one-stage
process at lower temperatures (600 °C). Hydrogen selective membranes are used to shift the
equilibrium towards hydrogen and to immediately produce ultrapure hydrogen (H2 purity
>99.99%).
In order to perform autothermal reforming at 600 °C, new catalyst and membranes are developed.
Pd-based hydrogen selective membranes were developed with an active layer of 4.5 µm Pd0.77Ag0.23
on a tubular ZrO2 support. These membranes have been tested and its performance has been
evaluated and compared to other known Pd-based membranes in literature. The comparison shows
that the newly developed membranes are performing better than most membranes in literature in
terms of hydrogen permeation. Membranes in literature which performed better in terms of
hydrogen permeation showed stability issues, whereas the membrane tested in this work, did not
show stability issues on the active layer.
A new 2 wt% Ru/CeZrO2 catalyst was developed and tested for autothermal reforming at 600°C in
a fluidized bed reactor. The catalyst showed mechanical stability but activity in terms of methane
conversion showed a decrease in activity over time. Next to the new catalyst, a reproducibility study
on 1.4 wt% Rh/ZrO2 was performed and showed similar results.
Catalyst and membranes were integrated in two reactor concepts. The packed bed membrane
microreactor (PBMMR) and the fluidized bed membrane reactor (FBMR). The PBMMR showed very
low methane conversions (>5%) which was not expected. The FBMR showed high methane
conversions up to 90% at 600 °C. Producing ultrapure hydrogen was not achieved due to the purity
of the extracted hydrogen. Further investigation on the membrane showed leakage at the sealing of
the membrane which is most probably caused by the temperature.
A theoretical comparison has been made between the two reactor concepts, considering simulation
results and an preliminary cost analysis. Simulation results showed the performance of both reactors
are similar in terms of methane conversion and hydrogen recovery. The temperature profiles of both
reactor concepts were different and the FBMR has better heat transfer within the bed. A preliminary
cost analysis has been made considering a small scale hydrogen production plant. The FBMR
concept had lower purchasing cost than the PBMMR. Based on both the experimental and the
theoretical comparison, the FBMR looks more promising for the autothermal reforming of methane
than the PBMMR concept.
Technische Universiteit Eindhoven University of Technology
II Table of Contents
Table of Contents
Abstract I
Table of Contents II
List of Figures IV
List of Tables VII
1. Introduction 1
2. Palladium based membranes 3 2.1.1 Flux expression for Pd-based membranes 3 2.2 Experimental procedure 5 2.2.1 Setup 5 2.2.2 External gas phase mass transfer limitations 6 2.2.3 CO poisoning effect 7 2.3 Results 8 2.3.1 ENEA self-supported planar Pd-Ag membrane 8 2.3.2 Tecnalia tubular Pd-Ag/ZrO2 membrane 9 2.4 Comparison with literature 12 2.5 Conclusion 14
3. Novel catalyst for ATR of methane 15 3.1 Catalyst properties 16 3.1.1 Particle size distribution 16 3.1.2 Carbon deposition 17 3.1.3 Determination of minimal fluidization velocity 18 3.1.4 Fluidization behavior 24 3.2 Catalyst stability 26 3.2.1 Experimental 26 3.2.2 Limitations of the system 27 3.2.3 Results 29 3.3 Conclusion 34
4. Evaluation of reactor concepts for ATR 35 4.1 Microreactor 36 4.1.1 Theory 36 4.1.2 Preparation and procedure 36 4.1.3 Results 37 4.1.4 Conclusion 39 4.2 Fluidized bed membrane reactor 40 4.2.1 Theory 40 4.2.2 Preparation and procedure 42 4.2.3 Results 45 4.3 Membrane stability in fluidized beds 49 4.3.1 Preparation of sealing protections 49 4.3.2 Results 49 4.4 Conclusion 55
5. Theoretical comparison reactor types for ATR of methane 56 5.1 Reactor concepts in literature 57 5.2 Modelling 58 5.2.1 Packed bed membrane microreactor 58 5.2.2 Fluidized bed membrane reactor 60 5.2.3 Comparison reactor types 63
Technische Universiteit Eindhoven University of Technology
III Table of Contents
5.3 Economical 65 5.4 Conclusion 67
6. Conclusion 68
7. Recommendations 69
Acknowledgement 70
Biblography 71
Appendix 74 A. Comparison of umf of CeZrO2 with literature 74 B. Parameters used in the packed bed reactor tests 75 C. Graphs FBMR with Tecnalia membranes 76 D. Preliminary cost analysis 77
Technische Universiteit Eindhoven University of Technology
IV List of Figures
List of Figures
Figure 1.1 Schematic overview of the ReforCELL project (“www.reforcell.eu,” 2013) .......................... 1 Figure 2.1 Schematic overview of the experimental setup for membrane testing ............................... 5 Figure 2.2 Sealing of the tubular Pd-Ag/ZrO2 membranes provided by Tecnalia, a schematic of the sealing (l) and one of the sealed membranes (with an additional protection ring) (r) ........................ 5 Figure 2.3 Effect of hydrogen dilution on the flux of hydrogen through the membrane (ENEA) at different transmembrane hydrogen partial pressures. The flux has been normalized to the case of 100% H2. .................................................................................................................................................. 6 Figure 2.4 The normalized flux of hydrogen through the membrane (ENEA) at 484 °C and feed flows of 400 to 470 Nml/min. The flux has been normalized to the case of 100% H2 and compared at different hydrogen partial pressures. ................................................................................................. 7 Figure 2.5 ENEA self-supported planar Pd-Ag membrane: averaged hydrogen flux as a function of the hydrogen transmembrane pressure difference at different temperatures (points: measurements; lines: predictions) ......................................................................................................... 8 Figure 2.6 Tecnalia tubular supported Pd-Ag membrane: averaged hydrogen flux as a function of the hydrogen transmembrane pressure difference at different temperatures (points: measurements; lines: predictions) ......................................................................................................... 9 Figure 2.7 Results for long-term hydrogen exposure to palladium membranes on different supports (Okazaki et al., 2009) ............................................................................................................ 10 Figure 2.8 Normalized hydrogen flow through the membrane over time. The results were obtained at 610 °C with a 100 vol% H2 stream during day and a 5 vol% H2 stream overnight. ...................... 10 Figure 2.9 Nitrogen leakage over time at 600 °C of the Pd-Ag/ZrO2 membrane developed by Tecnalia with the Swagelok graphite sealings ...................................................................................... 11 Figure 2.10 Predicted fluxes of several membranes at 600 °C at varying transmembrane hydrogen pressure ................................................................................................................................................... 13 Figure 2.11 Predicted fluxes of several membranes at 300 °C at varying transmembrane hydrogen pressure ................................................................................................................................................... 13 Figure 3.1 Particle size distribution of CeZrO2 after fluidization for 24 hours without (black) and after (red) sieving (125-250 µm fraction) .............................................................................................. 16 Figure 3.2 Particle size distribution of CeZrO2 after fluidization for 24 hours at 600 °C determined by dry sieving (sieves used: 100, 150, 212, 300 µm) ............................................................................. 17 Figure 3.3 The weight difference in terms of percentage of the weight of the sample over time during (1) activation: H2:N2 = 20:80, (2) Reforming reaction: CH4:H2O:N2 = 23:46:31 and (3) oxidation with air. Sample weights: CeZrO2 144.1 mg, Ru/CeZrO2 150.1 mg and Ni/Al2O3 46.9 mg. ................................................................................................................................................................ 18 Figure 3.4 example of Δp versus u0 for uniformly sized sharp sand. Umf is determined as the intersection between the fixed bed line with the horizontal line W/A (Kunii and Levenspiel, 1991). ................................................................................................................................................................ 19 Figure 3.5 Schematic of the minimal fluidization setup ..................................................................... 20 Figure 3.6 Δp versus u0 for the case of CeZrO2 particles with particle size 90-125 µm at 20 °C (l). Because of the wide distribution of particle size, the minimal fluidization velocity is determined as the intersection between the fixed bed line with the horizontal line W/At (r). ...................................21 Figure 3.7 The determined minimal fluidization velocity versus the temperature for several fractions of CeZrO2. .............................................................................................................................. 22 Figure 3.8 The determined minimal fluidization velocity versus the average particle size of the fractions of CeZrO2. .............................................................................................................................. 22 Figure 3.9 Experimental value for umf in comparison with theoretical models for different particle size distributions ................................................................................................................................... 23 Figure 3.10 Schematic of the experimental setup to determine the catalyst stability in a packed bed reactor of quartz ..................................................................................................................................... 26 Figure 3.11 Schematic of the quartz u-shaped reactor at the height of the catalyst and sand oven (l) actual reactor filled with diluted Ru/CeZrO2 (bed height = ±1 cm) (r) ............................................... 26 Figure 3.12 Methane conversion as a result of ATR of methane using a Rh/ZrO2 ........................... 29 Figure 3.13 Outlet composition on the stability test on CeZrO2 ......................................................... 30 Figure 3.14 Outlet composition (l) and methane conversion (r) on the stability test on Ru/CeZrO2
................................................................................................................................................................ 30 Figure 3.15 Temperature profile over time in the second batch stability test on Ru/CeZrO2 ........... 32
Technische Universiteit Eindhoven University of Technology
V List of Figures
Figure 3.16 Stable hydrogen production as found by Hybrid catalysis with ATR on the Ru/CeZrO2 catalyst (green line) ................................................................................................................................ 32 Figure 3.17 Outlet composition (l) and methane conversion (r) on the stability test on Ru/CeZrO2 with a WHSV of 22 h-1 ............................................................................................................................ 33 Figure 4.1 Schematic of a microreactor (A. L. Mejdell et al., 2009a) ................................................ 37 Figure 4.2 outlet composition of the ATR in a PBMMR with varying temperatures and pressures 38 Figure 4.3 Schematic representation of fluidized beds in different regimes (Kunii and Levenspiel, 1991) ....................................................................................................................................................... 40 Figure 4.4 Diagram of the Geldart classification of particles (Geldart, 1973) .................................... 41 Figure 4.5 Schematic overview of the setup in which the FBMR is tested ........................................ 43 Figure 4.6 schematic of the inside of the fluidized bed membrane reactor with five tubular membranes installed via the Swagelok graphite ferrules sealing method ......................................... 43 Figure 4.7 Autothermal reforming (ATR) in a FBMR with REB commercial hydrogen membranes at 600 °C at 1.3 and 1.47 bar. Feed: SCR of 1.91, OCR of 0.43 and an NCR of 5.24 with a total flow rate of 11.25 Nl/min. .............................................................................................................................. 45 Figure 4.8 Autothermal reforming (ATR) in a FBMR with REB commercial hydrogen membranes at 600 °C at 1.48 and 1.58 bar. Feed: SCR of 1.49, OCR of 0.43 and an NCR of 4.44 with a total flow rate of 12.4 Nl/min and SCR of 1.91, OCR of 0.43 and an NCR of 5.3 with a total flow rate of 11.25 Nl/min. ..........................................................................................................................................46 Figure 4.9 Two membranes with paste at the sealings to prevent particles from interacting with the graphite gasket, black insulation paste (l) and Ceramabond by Aremco (r) ......................................49 Figure 4.10 Nitrogen leakage and temperature of the reactor over time for conventional ferrule placement (top left), reversed ferrule placement (top right) and reversed ferrules placement + extra protection ring (bottom) ........................................................................................................................ 50 Figure 4.11 Normalized nitrogen leakage after 500 °C for the conventional placed ferrules without fluidization conditions, reversed and reversed + protected ferrule placement under fluidization conditions ............................................................................................................................................... 50 Figure 4.12 Nitrogen leakage after exposing the membrane to several gases at different temperatures and durations ................................................................................................................... 51 Figure 4.13 image of a membrane used in the FBMR tested in water with air at 1 barg on the inside of the membrane ................................................................................................................................... 52 Figure 4.14 membrane with the depicted areas analyzed by SEM-EDX ............................................ 52 Figure 4.15 SEM pictures of a membrane classified as having a high amount of particles on the surface at the central area (a) (top left and right) the interphase between membrane and graphite (c) (bottom left) and the graphite zone (d) (bottom right) .................................................................... 53 Figure 4.16 mapping surface of the central zone of a membrane classified as having a high amount of particles by EDX ................................................................................................................................. 53 Figure 4.17 SEM image (l) and EDX analysis (r) on the membrane classified as high surface roughness and some sheets .................................................................................................................. 54 Figure 5.1 Schematic representation of the PBMMR model ............................................................... 58 Figure 5.2 comparison simulation results between the base case without membrane integration and with the isothermal case ................................................................................................................ 59 Figure 5.3 Comparison of different SCR ratios without membrane integration (l) and with membrane integration (r) ..................................................................................................................... 59 Figure 5.4 Comparison methane conversion in different systems (l) and the temperature profile of the adiabatic simulation (r) .................................................................................................................. 60 Figure 5.5 Comparison methane conversion with cooling (l) and the accompanying temperature profiles (r) .............................................................................................................................................. 60 Figure 5.6 Schematic representation of the FBMR model (Patil, 2005) ........................................... 60 Figure 5.7 comparison methane conversion in the FBMR ................................................................. 61 Figure 5.8 Comparison methane conversion with different catalyst loadings in the bed ................. 62 Figure 5.9 Comparison methane conversion with different specific membrane areas .................... 62 Figure 5.10 Comparison of the FBMR and PBMMR model in terms of methane conversion and hydrogen permeation ............................................................................................................................ 63 Figure 5.11 Representation of fluidized bed (l) and the microreactor (r) ............................................ 65 Figure 1 parity plots comparing the experimentally determined minimal fluidization velocity with the predicted values according to four different correlations ............................................................. 74 Figure 2 Steam methane reforming (SMR) in a FBMR with five tubular Tecnalia membranes at 500 and 550 °C and a pressure of 1.3 bar. Feed: SCR of 3 and an NCR of 8.4 with a total flow rate of 10.3 Nl/min. ........................................................................................................................................... 76
Technische Universiteit Eindhoven University of Technology
VI List of Figures
Figure 3 Steam methane reforming (SMR) in a FBMR with five membranes at 550 and 600 °C. Feed: SCR of 3 and an NCR of 8.4 with a total flow rate of 10.3 Nl/min to ensure fluidization. ..... 76 Figure 4 Autothermal reforming of methane (ATR) in a FBMR with five membranes at 600 °C. Feed: SCR of 3, OCR of 0.25 and an NCR of 8.2 with a total flow rate of 10.3 Nl/min to ensure fluidization. ............................................................................................................................................ 76
Technische Universiteit Eindhoven University of Technology
VII List of Tables
List of Tables
Table 2.1 Summary of published data for hydrogen permeability of palladium based membranes .12 Table 3.1 minimal fluidization velocity for several particle size distributions of CeZrO2 ..................21 Table 3.2 Values of several investigators of the two constants in equation (Kunii and Levenspiel, 1991) ....................................................................................................................................................... 23 Table 3.3 Coefficient of determination (R-squared) for the observed data and predicted model values versus the fitted line y=x ............................................................................................................ 24 Table 3.4 Composition of the filler used for the FBMR, developed by Hybrid Catalysis .................. 24 Table 3.5 Pictures of the segregation test of CeZrO2 (yellow) with the zirconia based filler particles (white) in a 2D fluidized bed ................................................................................................................ 25 Table 3.6 BET analysis on the rhodium and ruthenium based catalyst and supports ....................... 31 Table 4.1 Summary of the ATR test in a packed bed membrane microreactor with a SCR of 1.9, OCR of 0.44 and NCR of 3 at different temperatures and pressures in- and excluding separation with vacuum on the permeate side....................................................................................................... 38 Table 4.2 Summary of the FBMR test performed with the REB membranes at 600 °C at 1.3 and 1.47 bar ...................................................................................................................................................46 Table 4.3 Summary of the FBMR test performed with the REB membranes at 600 °C at 1.48 and 1.58 bar with different feed flowrates. ..................................................................................................46 Table 4.4 Summary of SMR in the FBMR with Tecnalia tubular Pd-Ag/ZrO2 membranes at 1.3 bar at several temperatures. ........................................................................................................................ 47 Table 4.5 Comparison of ATR/SMR in the FBMR with REB and Tecnalia membranes at 600 °C. 48 Table 4.6 Classification given to membranes based on optical microscopy ...................................... 52 Table 4.7 Summary of the EDX results of the membrane with high amount of particles on the surface (composition in wt%) ............................................................................................................... 54 Table 5.1 parameters and constants used for the modelling of ATR in a PBMMR ........................... 58 Table 5.2 parameters and constants used for the modelling of ATR in a FBMR .............................. 61 Table 5.3 Effect of catalyst loading on several indicators for ATR ...................................................... 62 Table 5.4 Effect of specific membrane area on several indicators for ATR ........................................ 62 Table 5.5 parameters and constants used for the modelling of ATR in a PBMMR and FBMR ........ 63 Table 5.6 Results of the simulations of the PBMMR and FBMR ....................................................... 63 Table 5.7 Preliminary cost analysis for the membrane assisted reactor concepts for a small scale production plant ................................................................................................................................... 66 Table 1 Coefficient of determination (R-squared) for the observed data and predicted model values versus the fitted line y=x ....................................................................................................................... 74 Table 2 Parameters as used in the stability test for ATR of methane in a packed bed reactor ......... 75
Technische Universiteit Eindhoven University of Technology
1 Introduction
1. Introduction
In recent years, concerns have been growing worldwide regarding the environmental consequences
of heavy dependence on fossil fuels, particularly climate change. Concerns rise about the security of
energy supply, increased fuel prices and carbon emissions. At the present rate of consumption, the
known petroleum resources are expected to be depleted in less than 50 years. Therefore new energy
sources are being searched for the long term, especially the interest in bio-fuel has grown strongly
in recent years (Balat and Balat, 2009). For the short- to mid-term future hydrogen is a very
interesting candidate to contribute to the energy demand (Arzamendi et al., 2013).
Hydrogen is used in many industries and is one of the main commodities in the chemical sector.
Regarding the energy system, hydrogen offers significant advantages. It can be used for almost every
sector which requires energy (transport, households, industry etc.)(Holladay et al., 2009). When
used, the only product is water vapor, making it very friendly for the end-user. Moreover, hydrogen
can be produced from a wide variety of resources (Muellerlanger et al., 2007).
Hydrogen production is mostly done by using fossil fuels, such as natural gas and coal. Alternative
energy sources such as nuclear, solar, wind and biomass are interesting as they are capable of
producing hydrogen with a sustainable fuel cycle (no carbon emission during production and end
use) (Bartels et al., 2010). But before this technology is market ready, a lot of progress can still be
made in the conventional production methods.
Approximately 96% of the hydrogen produced is from fossil fuel-based processes of which 48%
from natural gas, 30% from oil and 18% from coal (Balat and Balat, 2009; Kothari et al., 2008). The
current main production processes for the fossil fuel-based hydrogen are natural gas steam
reforming, partial oxidation of hydrocarbons and coal gasification. The natural gas steam reforming
is the most used process and therefore it is interesting to improve this process particularly as a lot of
energy saving can still be achieved. The ReforCell project is aiming to the optimization of ultrapure
hydrogen production.
Figure 1.1 Schematic overview of the ReforCELL project (“www.reforcell.eu,” 2013)
The main project aims are to develop a micro combined heat and power (m-CHP) system by
developing a novel, more efficient and cheaper multi-fuel membrane reformer for pure hydrogen
production. Intensification of the process is tried to be reached by integration of reforming and
purification in one single step. Furthermore the project aims at optimizing the energy losses in the
system by optimizing the design of heat exchangers and circumventing the mass and heat transfer.
Technische Universiteit Eindhoven University of Technology
2 Introduction
This work is focused on the development of the membrane reformer including the catalyst,
membranes and reactor design.
Steam reforming of methane is a process which at industrial scale is done in three reaction steps to
produce hydrogen and remove the CO. First, methane and steam are converted to hydrogen and CO
by the steam methane reforming, Eq. 1.1, using a (mostly nickel) catalyst. Next is the water gas shift
which converts most of the CO in the less dangerous CO2 and also converts steam into the desired
hydrogen, Eq. 1.2. The third step consists of removal of any excess CO by preferential oxidation, Eq.
1.3.
4 2 23catCH H O H CO ( 0)H 1.1
2 2 2catH O CO H CO ( 0)H 1.2
2 22 2CO O CO ( 0)H 1.3
For this current process, at least four reactors are needed as these steps take place in separate
reactors, WGS even in two separate reactors. Another method exist in which SMR, WGS and the
oxidation of methane, Eq. 1.4, are combined in one reactor: auto thermal reforming of methane
(ATR), Eq. 1.5. Autothermal condition requires the feed of inlet gases in a way, that the overall heat
production equals zero. This can be achieved by optimizing the oxygen to carbon ratio (OCR or ).
4 2 2 22CH O CO H O ( 0)H 1.4
4 2 2 2 22(1 ) 2(2 )catCH O H O CO H 1.5
The advantage of this process is that only one reactor is needed and excess steps are eliminated. One
drawback of this process is the limitation in the thermodynamic equilibrium. Therefore hydrogen
yields are lower than achieved in the industrial applied process. A way to improve the overall yield is
extracting hydrogen during the reaction. This can be achieved by placing hydrogen selective
membranes in the reactor. The two main advantages using membranes in this process are a higher
overall yield and the production of a pure hydrogen product stream.
Current ATR processes are running at temperatures between 800 to 900 °C, which is mainly due
to limitations in the thermodynamic equilibrium of the system. In this work, new catalyst and
hydrogen selective membranes are integrated in one system to run the ATR reaction at lower
temperature (600 °C) and maintain a high conversion and yield.
Because of the lower operating temperature a catalyst has to be found which is able to operate at the
required conditions. Therefore two catalysts Rh/ZrO2 and Ru/Ce0.25Zr0.75O2 have been evaluated.
The performance of this system will be evaluated in two different reactors, the membrane
microreactor and fluidized bed membrane reactor.
The ultimate goal is to determine the best operating system for autothermal reforming of methane.
Including reactor type, catalyst and membrane integration.
Technische Universiteit Eindhoven University of Technology
3 Palladium based membranes
2. Palladium based membranes
To enhance the yield of the ATR reaction, a hydrogen selective membrane is required to selectively
extract hydrogen from the system. Palladium based membranes allow proton diffusion through the
metal surface and are therefore selective to hydrogen. Silver is added (23 wt%) to reduce the critical
temperature for β-embrittlement of the membrane which normally occurs below 300 °C (Gallucci et
al., 2013).
Two membranes have been investigated in this work. A planar self-supported Pd-Ag membrane
developed by ENEA and a tubular Pd-Ag/ZrO2 membrane developed by Tecnalia. In this chapter,
the permeation or flux through the membrane will be evaluated. Next to the permeation it is essential
for a membrane to be selective. The selectivity of the membrane for hydrogen compared to nitrogen
will be evaluated in this chapter, whereas the selectivity towards other gases like CO is discussed in
Chapter 4.
2.1.1 Flux expression for Pd-based membranes
Hydrogen permeation through membranes can be defined as five steps:
1) Diffusion from the gas phase to the metal surface on the feed side
2) Adsorption on the metallic surface and dissociation into H atoms
3) Diffusion through the metal lattice as protons
4) Regeneration of H atoms into a H2 molecule and desorption from the metal surface
5) Diffusion from the metal surface into the permeate gas phase
The first step can be considered as external mass transfer limitation which can be avoided by using
pure hydrogen as feed stream or maintaining a constant hydrogen partial pressure at the feed side
by increasing the total hydrogen flow. The combination of the steps can be rate-controlling for the
overall process. The hydrogen flux through the membrane, JH2,can be expressed into an expression
commonly known as Sieverts’ law:
2 ,ret ,perm2 2H H
n nH
PJ p pt
3.1
where P is the membrane permeability [mol/s/m/Pan], t is the membrane thickness [m], 2H
p the
hydrogen partial pressure in either the retentate or permeate side. The value of n depends on the
rate limiting steps of the permeation process. If membrane bulk diffusion is rate limiting, the value
of n is approximately equal to 0.5. The permeance P
t is typically described with an Arrhenius type
dependency on the temperature:
0 exp actEP kt RT
3.2
with 0k the hydrogen permeability constant [mol/s/m2/Pan], R the gas constant [J/mol/K], T the
temperature [K] and Eact activation energy for membrane permeability [J/mol].
Technische Universiteit Eindhoven University of Technology
4 Palladium based membranes
The flux through the membrane is dependent on several factors next to the hydrogen partial pressure
difference over the membrane like the permeability and the thickness of the membrane. External
limitations can reduce the flux including concentration polarization and CO poisoning.
Concentration polarization occurs if the flux of a species (in this case hydrogen) is higher than its
flux in the bulk. This causes a decrease in concentration of hydrogen in the bulk-membrane interface
and an increase in concentration of hydrogen on the permeate side. This effects strongly reduces the
driving force (the hydrogen transmembrane pressure difference) at a local level. Concentration
polarization can be prevented by increasing the mixing in the bulk phase for instance by placing
baffles, which however could result in higher resistances and larger pressure drops over the reactor.
Another known limitation is the poisoning effect of CO. When CO is present in the bulk, depending
on concentration and temperature, hydrogen permeation is reduced. The most accepted explanation
for this behavior is that CO adsorbs to the Pd-Ag surface and therewith blocking available sites for
hydrogen to adsorb and dissociate in hydrogen atoms (Gallucci et al., 2007; a. L. Mejdell et al., 2009;
Miguel et al., 2012).
Technische Universiteit Eindhoven University of Technology
5 Palladium based membranes
2.2 Experimental procedure
2.2.1 Setup
The setup that has been used to determine the flux expression is designed specifically for the
measurement of the permeability of H2 trough a membrane.
Figure 2.1 Schematic overview of the experimental setup for membrane testing
The setup allows to feed gases as mixtures to the reactor. The membrane is tested in a shell and tube
configuration of reactor, whereas the pressure in the reactor can be manipulated by a back pressure
regulator in the retentate stream. The flow on the permeate side was measured using a Horibastec
Film Flow meter.
2.2.1.1 Sealing of the Tecnalia tubular Pd-Ag/ZrO2 membranes
The membranes provided by Tecnalia had to be sealed in order to be usable at high temperatures.
For this purpose, a sealing method was developed which consists of a graphite gasket and a Swagelok
connection (Chen et al., 2010). This sealing method was further investigated and improved within
the SMR group.
Figure 2.2 Sealing of the tubular Pd-Ag/ZrO2 membranes provided by Tecnalia, a schematic of the sealing (l) and one of the
sealed membranes (with an additional protection ring) (r)
The graphite gasket is used instead of a standard stainless steel ferrule, because it is softer and can
be compressed without damaging the membrane.
H2
CO2
N2
CO
F
F
F
F
PI
Feed Section
PCV
PI
Vent
GC
Vent
GC
Bypass
TI
Retentate
Permeate
PI
PCV
Technische Universiteit Eindhoven University of Technology
6 Palladium based membranes
At higher temperatures, the graphite will soften more and will also expand, thus improving the
sealing. From experience it was discovered that retightening the Swagelok connection was beneficial
in terms of leakage. Retightening too much, however, resulted in breaking the porous ZrO2 support
of the membrane.
During the sealing and resealing process, the sealings were tightened using a torque controlled
wrench. After some sealing processes, the maximum torque was decided to be around 8 Nm which
corresponds to about 1 Nml/min N2 at a pressure difference of 4 bar. At elevated temperatures, this
leakage would get less due to the expansion of the graphite gasket and would result in the desired
ultrapure hydrogen.
Before integration of the membranes into a reactor, preliminary leaking tests were conducted. A
visual test where 1 barg of either air or helium was used in the tube side of the membrane. The
membrane was then placed in a liquid (either water or ethanol) to check for bubbles.
Another preliminary test was done by installing the membrane in a tube which could be pressurized.
A pressure of 4 barg was applied to the membrane and the permeate flow was be measured.
2.2.2 External gas phase mass transfer limitations
To obtain an accurate description of the hydrogen flux through the membrane, it should be
confirmed that gas phase mass transfer limitations are negligible. This was done by feeding different
compositions of N2 and H2, keeping the transmembrane hydrogen pressure the same, but differing
the hydrogen concentration, and comparing the flux to the partial pressure difference of hydrogen.
A few compositions have been evaluated to verify the measurements are free of any external gas
phase mass transfer limitations. The results are shown in figure 2.3 and it can be seen that the effect
of diluting the hydrogen is negligible for the ENEA membrane.
Figure 2.3 Effect of hydrogen dilution on the flux of hydrogen through the membrane (ENEA) at different transmembrane hydrogen
partial pressures. The flux has been normalized to the case of 100% H2.
From Figure 2.3 it can be concluded that mass transfer limitations can be neglected until 50 vol%
H2 in the feed as the permeation of hydrogen stays the same. For the determination of membrane
constants, external gas phase mass transfer limitation do not occur as these tests are conducted in a
pure H2 atmosphere, both in the retentate and permeate side.
25 50 75 1000.80
0.85
0.90
0.95
1.00
H2 transmembranepressure (bar)
0.5 1.0 1.5 2.0
Nor
mal
ized
hyd
roge
n flu
x (-)
H2/N2 mixtures (vol% H2)
498 °C
Technische Universiteit Eindhoven University of Technology
7 Palladium based membranes
When performing experiments with ATR of methane, the hydrogen concentration in the reactor will
not be above 50% and external mass transfer limitations should be taken into account. Furthermore
it should be taken into account that for the fluidized bed membrane reactor a pressure drop is present
over the length of the bed increasing the effect of concentration polarization.
2.2.3 CO poisoning effect
During the SMR reaction, CO is produced and can reduce the permeability of the palladium-based
membrane. To verify this effect, H2/CO mixtures were fed to the ENEA membrane and the H2 flux
was measured. The hydrogen transmembrane pressure was kept constant.
Figure 2.4 The normalized flux of hydrogen through the membrane (ENEA) at 484 °C and feed flows of 400 to 470 Nml/min.
The flux has been normalized to the case of 100% H2 and compared at different hydrogen partial pressures.
From Figure 2.4 it can be seen that there is an effect of CO on the hydrogen flux through the
membrane. The CO concentration during ATR is expected to be around 4 vol% so the effect of CO
is almost negligible. Studies done by Mejdell and Miguel show that with rising temperature, the CO
poisoning effect on Pd-Ag membranes gets smaller ( a. L. Mejdell et al., 2009; Miguel et al., 2012).
ATR will be performed at 600 °C, therefore it is expected that the CO poisoning effect is negligible
due to the higher temperature and lower CO concentration.
0 5 10 15
0.90
0.95
1.00H2 transmembrane pressure (bar)
0.5 1.0 1.5 2.0 2.5
Nor
mal
ized
hyd
roge
n flu
x (-)
H2/CO mixtures (vol% CO)
484 °C
Technische Universiteit Eindhoven University of Technology
8 Palladium based membranes
2.3 Results
The results are presented in two parts. The hydrogen permeability and the membrane parameters
has been determined for the ENEA self-supported planar Pd-Ag membrane in the first part and for
the Tecnalia Pd-Ag/ZrO2 tubular membrane in the second part. Additionally the effect of the support
on the active Pd-Ag layer in terms of hydrogen permeation has been determined for the Pd-Ag/ZrO2
membrane.
2.3.1 ENEA self-supported planar Pd-Ag membrane
The first membrane is a planar self-supported Pd-Ag membrane developed by ENEA. Experiments
at different pressures and temperatures were used to evaluate the hydrogen flux through the
membrane. pure hydrogen was used that no external gas transfer limitations can limit the
permeation of hydrogen.
The membrane was tested in a microreactor with 6 channels of the dimensions 1 x 1 x 13 mm,
resulting in an effective membrane area of 7.8·10-5 m2.
Figure 2.5 ENEA self-supported planar Pd-Ag membrane: averaged hydrogen flux as a function of the hydrogen transmembrane
pressure difference at different temperatures (points: measurements; lines: predictions)
From this data, the following permeability constant and activation energy have been found:
Permeability constant k0 2.58E-07 mol/(m2·Pa0.5·s)
Eact 10.45 kJ/mol
The value of n was determined to be 0.5. Performing an error analysis, the best linear fit was found
for n=0,5, with the restriction, that the fit should pass the origin.
For this membrane, the nitrogen leakage was unmeasurable, because the Film flow meter can only
detect flows >0.2ml/min. Even at high pressure differences, no flow could be detected.
In chapter 2.4 the membrane will be compared to other known Pd-Ag membranes.
0 50 100 150 200 250 3000.0
0.1
0.2
0.3
0.4
Temperature (°C) 470 520 576 630A
vera
ge h
ydro
gen
flux
(mol
s-1 m
-2)
Transmembrane pressure difference p0.5H2ret-P
0.5H2perm (Pa0.5)
Technische Universiteit Eindhoven University of Technology
9 Palladium based membranes
2.3.2 Tecnalia tubular Pd-Ag/ZrO2 membrane
Tecnalia developed a tubular Pd-Ag on ZrO2 membrane, designed for the use in a fluidized bed. The
thickness of the active Pd-Ag layer was about 4.5 µm. The tube had a diameter of 10.3 mm and a
length of 42 mm , resulting in an effective membrane area of 1.36·10-3 m2. The membrane was sealed
with graphite gaskets in Swagelok connections as described in chapter 2.2.1.1.
2.3.2.1 Permeability
The hydrogen permeability was measured using the same method as for the ENEA membrane.
Figure 2.6 Tecnalia tubular supported Pd-Ag membrane: averaged hydrogen flux as a function of the hydrogen transmembrane
pressure difference at different temperatures (points: measurements; lines: predictions)
From this data, the following permeability constant and activation energy have been found:
Permeability constant k0 6.93E-08 mol/(m2·Pa0.5·s)
Eact 9.99 kJ/mol
Also in this case, the value for n= 0.5 gives the best linear fit to the experimental data, which will
later on be used to predict the hydrogen recovery in the fluidized bed membrane reactor. The results
will also be compared to hydrogen membranes known in literature in chapter 2.4.
The nitrogen leakage for this membrane was also evaluated and these results can be found in the
next chapter (2.3.2.2).
2.3.2.2 Long-term effects
The interaction between the Palladium and the chosen support proved to be important as found by
(Okazaki et al., 2009). Palladium on alumina and palladium on yttrium-stabilized zirconia (YSZ)
have been investigated and the palladium on alumina showed a steep decrease in hydrogen
permeation within a few days whereas the palladium on YSZ is going to a stable value.
0 50 100 150 200 2500.0
0.1
0.2
0.3
0.4
0.5
0.6
0.7
0.8
Temperature (°C) 385 455 490 545 598A
vera
ge h
ydro
gen
flux
(mol
s-1 m
-2)
Transmembrane pressure difference p0.5H2ret-p
0.5H2perm (Pa0.5)
Technische Universiteit Eindhoven University of Technology
10 Palladium based membranes
Figure 2.7 Results for long-term hydrogen exposure to palladium membranes on different supports (Okazaki et al., 2009)
It is important to know the effect of the support on the performance of the membrane. Therefore the membrane has been exposed to a 5 vol% H2 feed at 600 °C. The reactor was kept at 1.6 bar during the day and for safety reasons at 1 bar at night. Hydrogen permeation has been recorded and can be found in figure 2.8.
Figure 2.8 Normalized hydrogen flow through the membrane over time. The results were obtained at 610 °C with a 100 vol% H2
stream during day and a 5 vol% H2 stream overnight.
The membrane shows the same trend as Palladium/YSZ as found by Okazaki et al. where the normalized flow stabilized at 90%. In terms of hydrogen flux through the membrane, no problems are expected using the membrane in the fluidized bed application. Next to hydrogen permeation the hydrogen selectivity is important. Therefore the nitrogen leakage was measured during this long-term test.
Technische Universiteit Eindhoven University of Technology
11 Palladium based membranes
Figure 2.9 Nitrogen leakage over time at 600 °C of the Pd-Ag/ZrO2 membrane developed by Tecnalia with the Swagelok graphite
sealings
As can be seen in Figure 2.9 the nitrogen leakage increases over time and is 3.5 times higher after
five days compared to the start of the long-term experiment. This could be a problem using the
membranes in the fluidized bed membrane reactor to extract ultrapure hydrogen. To extract
ultrapure hydrogen, a high nitrogen leakage is not desired and should be monitored carefully during
the experiments.
The selectivity to hydrogen of the membrane can be expressed by the following term:
2,
2,
perm
perm
H
N
JS
J 3.3
The selectivity at the start of the long term experiment was 12000 and dropped to 1600 after 120
hours of exposing the membrane to a temperature of 600 °C, hydrogen and nitrogen as feed gases.
For producing ultrapure hydrogen a selectivity of at least 10000 is desired.
0 20 40 60 80 100 120
1.0
1.5
2.0
2.5
3.0
3.5
Nitr
ogen
leak
age
norm
aliz
ed (-
)
Time (h)
2.0x10-12
3.0x10-12
4.0x10-12
5.0x10-12
Nitr
ogen
leak
age
(mol
s-1 P
a-1)
Technische Universiteit Eindhoven University of Technology
12 Palladium based membranes
2.4 Comparison with literature
Now that the permeability for the membranes have been determined, it’s interesting to see how they
compare to other Pd-Ag membranes stated in literature in terms of hydrogen permeation. For this
several literature found membranes are compared in terms of hydrogen permeation to the
membranes in this work.
There are several researchers working on Pd-Ag membranes. Some membranes are chosen for the
comparison because of their geometry, others by their thickness and others by their novelty. (A. L.
Mejdell et al., 2009a; Miguel et al., 2012; Morreale et al., 2003; Vadrucci et al., 2013)
Table 2.1 Summary of published data for hydrogen permeability of palladium based membranes
Thickness (µm)
Geometry Composition T (K) k0 (mol/(m s Pan)) Eact (kJ/mol) n
Miguel 50 Tubular Pd0.75-Ag0.25 473-573 1.16 × 10−5 17.41 0.5
Morreale 1000 Planar Pd 623-1173 1.92 × 10−7 13.81 0.5
Vadrucci 84 Tubular Pd0.75-Ag0.25 473-623 2.95 × 10−8 2.531 0.5
Vadrucci 150 Tubular Pd0.75-Ag0.25 473-623 5.63 × 10−8 5.456 0.5
Vadrucci 200 Tubular Pd0.75-Ag0.25 473-623 2.06 × 10−8 2.592 0.5
Mejdell 1.4 Planar Pd0.77-Ag0.23 573 n.d. n.d. 0.5
Mejdell 2.2 Planar Pd0.77-Ag0.23 573 n.d. n.d. 0.5
REB 45 Tubular Pd - 1.70 × 10−10 6.170 0.72
ENEA (this work)
50 Planar Pd0.77-Ag0.23 723-873 2.58 × 10−7 10.45 0.5
Tecnalia (this work)
4.5 Tubular Pd0.77-Ag0.23 673-873 6.93 × 10−8 9.99 0.5
From this data it can be concluded that the activation energy and the permeability constant of both
membranes are within the same range as other reported palladium membranes. It is more
interesting to see the permeability versus the temperature.
Technische Universiteit Eindhoven University of Technology
13 Palladium based membranes
Figure 2.10 Predicted fluxes of several membranes at 600 °C at varying transmembrane hydrogen pressure
To compare the membranes the fluxes of the different membranes are predicated as function of
partial hydrogen pressure difference at later operating conditions, as can be seen in Figure 2.10. It
can be concluded that the Tecnalia tubular membrane with an active layer of Pd-Ag of 4.5 µm has
the best performance at 600 °C, whereas the ENEA membrane also showed better performance
compared to the membranes reported in literature.
Figure 2.11 Predicted fluxes of several membranes at 300 °C at varying transmembrane hydrogen pressure
If the predicted fluxes are compared at 300 °C (temperature at which the permeability of the SINTEF
membranes is defined) it can be seen that the SINTEF membranes are performing significantly
better than all other membranes. Although the reported permeability was similar to the membrane
of Tecnalia, the flux is three times higher due to the smaller thickness of the SINTEF membranes.
Although better hydrogen fluxes are found with the thinner membranes of SINTEF, it reported that
the membrane unfortunately had a decreasing selectivity to hydrogen over time and the formation
of pinholes and other defects were clearly visible (A. L. Mejdell et al., 2009a).
In summary, for the planar membrane of ENEA, when taking into account the thickness of the
membrane, the performance in terms of hydrogen flux is quite reasonable. The Tecnalia membrane
also showed higher hydrogen permeation than other reported membranes. For membranes with
higher hydrogen permeability, stability issues had been reported.
250 300 350 400 450 500 550 600 6500.0
0.5
1.0
1.5
2.0
2.5
Hyd
roge
n flu
x th
roug
h th
e m
embr
ane
(mol
s-1 m
-2)
Transmembrane H2 pressure difference P0.5H2ret-P
0.5H2perm (Pa0.5)
1000 µm Pd 200 µm REB 50 µm ENEA (this work) 4.5 µm Tecnalia (this work)
T = 600 °C
250 300 350 400 450 500 550 600 6500
1
2
3
4
5
6
7
8
9
T = 300 °C
Hyd
roge
n flu
x th
roug
h th
e m
embr
ane
(mol
s-1 m
-2)
Transmembrane H2 pressure difference P0.5H2ret-P
0.5H2perm (Pa0.5)
1000 µm Pd 84 µm ENEA 150 µm ENEA 200 µm ENEA 1.4 µm SINTEF 2.2 µm SINTEF 50 µm ENEA (this work) 4.5 µm Tecnalia (this work)
Technische Universiteit Eindhoven University of Technology
14 Palladium based membranes
2.5 Conclusion
Both the ENEA and Tecnalia membranes have been tested and the permeability constant and
activation energy have been determined. From the preliminary test it is concluded that external gas
transfer limitations should be taken into account at lower concentrations of hydrogen (< 50 vol%)
and the CO effect shows that in terms of hydrogen permeation, the inhibition is negligible at the
expected conditions during ATR of methane.
In comparison to published palladium membranes it can be concluded that both membranes
perform relatively well in terms of hydrogen permeation. The permeability constants are lower than
the published membranes but the thickness is much smaller resulting in excellent hydrogen flows.
Membranes have been published with a very thin Pd-Ag layer which perform better in terms of
hydrogen flux, these membranes however don’t show the required mechanical stability which is
desired to produce ultrapure hydrogen.
Technische Universiteit Eindhoven University of Technology
15 Novel catalyst for ATR of methane
3. Novel catalyst for ATR of methane
For conventional hydrogen production processes a nickel-based catalyst is often used, that works at
elevated temperatures (800 – 900 °C) and is therefore not suitable for the proposed ATR system. In
the ATR system, the WGS reaction takes place and these high temperatures are not favorable for the
thermodynamic equilibrium of this reaction. Pd-based membranes are used which show instability
at elevated temperatures. Therefore, a catalyst is needed for lower temperatures and Rh/ZrO2 and
Ru/CeZrO2 are promising candidates.
Rh/ZrO2 has already been tested by P. Wolbers and the data has been used in this work (Wolbers,
2013). The Ru/CeZrO2 is a novel catalyst and no research on this catalyst has been done before.
Experiments on the Ru/CeZrO2 catalyst include the mechanical stability of the particles, the
suitability for application in fluidized beds and the activity stability of the catalyst during autothermal
reforming of methane.
Technische Universiteit Eindhoven University of Technology
16 Novel catalyst for ATR of methane
3.1 Catalyst properties
In this work a 1.4 wt% Rh/ZrO2 catalyst developed by P. Wolbers at the TU/e has been used. A study
has already been performed on this catalyst and the results from that study will be used in this work.
This catalyst will only be discussed in chapter 3.2.3.1, results of the reproducibility test.
A new 2 wt% Ru/Ce0.75Zr0.25O2 catalyst developed by Hybrid Catalysis has also been used. It is
designed to perform ATR in a fluidized bed and therefore the experiments conducted on this catalyst
are to verify the viability of this catalyst to be used in a fluidized bed.
3.1.1 Particle size distribution
The particle size distribution of the provided catalyst is in the range of 125 to 250 µm (sieve fraction).
Because this catalyst is used in a fluidized bed system, it is important to know if the particles can
withstand the fluidization and check for negative effects like particles sintering, abrasion etc.
Because of the mentioned reasons the particle size distribution has been measured before and after
fluidization and at different temperatures.
To evaluate the particle size distribution, the CeZrO2 particles were fluidized for 24 hours and the
particle size distribution was determined using a Fritsch Analysette 22 MicroTec plus. This machine
analyzes a dispersion (water with the selected particles) and the intensity of light scattered when a
laser beam passes through a dispersed particulate sample is measured. From this data, the provided
software calculates the particle size distribution.
Figure 3.1 Particle size distribution of CeZrO2 after fluidization for 24 hours without (black) and after (red) sieving (125-250 µm
fraction)
It has been noticed that after fluidization the particle size distribution showed a high amount of
small particles which was not expected from optical comparison of the samples. Therefore the
particles were sieved and the 125-250 µm (the range that was supplied) was measured again. As can
be seen in Figure 3.1, there is no significant difference in measured particle size distribution between
the two samples.
The reason for the detected small particles is most probably, that the CeZrO2 particles, which are
mechanically pressed from powder, can fall apart in water in combination with the ultrasonic bath.
Also after sieving, no significant trace of small particles were found.
The temperature for which this catalyst is designed is to work at 600 °C, therefore the support has
also been fluidized for 24 hours at this temperature.
Technische Universiteit Eindhoven University of Technology
17 Novel catalyst for ATR of methane
Figure 3.2 Particle size distribution of CeZrO2 after fluidization for 24 hours at 600 °C determined by dry sieving (sieves used:
100, 150, 212, 300 µm)
From the dry sieving it can be concluded that the particle size distribution is in the same range and
therefore fluidization has no significant effect on the particle size distribution, making it suitable as
a material for a fluidized bed.
3.1.2 Carbon deposition
Before the catalyst will be implemented in the system, it’s important to determine the properties of
the catalyst. One of these properties is the carbon deposition as this causes a drop in activity of the
catalyst. To determine if the catalyst is prone to carbon deposition, the sample has been tested using
thermo gravimetric analysis (TGA).
3.1.2.1 Setup and procedure
The tests are performed in a TGA. Next to the catalyst, CeZrO2 was also tested to check if the support
accounts for carbon deposition. Ni/Al2O3 was tested as a reference material to verify the
comparability with previous obtained results from the high pressure TGA by P.F. Wolbers.
All experiments were performed in the same conditions and the same procedure has been used
which consists of three stages:
1. Activation of the catalyst at with a flow of 0.02 mol/min consisting of 20% hydrogen in
nitrogen at 700 °C for 2 hours.
2. Reforming reaction with a flow of 0.02 mol/min consisting of 23% methane, 46% steam
and 31% nitrogen (SCR = 2) at 700 °C for 15 hours.
3. Oxidation of the catalyst with a flow of 0.02 mol/min of air at 700 °C for 2 hours.
For the CeZrO2 and Ru/CeZrO2 around 150 mg was used for the experiment. For the Ni/Al2O3 an
amount of 50 mg was used and the reforming reaction was performed for 5 hours.
3.1.2.2 Results
Graphs of the results are plotted in Figure 3.3. It can be seen that there is no significant weight
increase during the reforming reaction. Also no visual changes of the sample were observed after
the experiment.
0.0
10.0
20.0
30.0
40.0
50.0
60.0
<75 125 181 256 >350P
erc
en
tage
(%
)
Particle size (µm)
Technische Universiteit Eindhoven University of Technology
18 Novel catalyst for ATR of methane
Figure 3.3 The weight difference in terms of percentage of the weight of the sample over time during (1) activation: H2:N2 = 20:80,
(2) Reforming reaction: CH4:H2O:N2 = 23:46:31 and (3) oxidation with air. Sample weights: CeZrO2 144.1 mg, Ru/CeZrO2 150.1 mg
and Ni/Al2O3 46.9 mg.
Comparing the support and the Ruthenium catalyst it can be seen that the behavior is similar,
indicating that carbon deposition is not an issue for the Ruthenium catalyst during 15 hours of
reforming reaction. In the stability test, which will be discussed in chapter 3.2.3.2, a drop in catalyst
activity was observed. Oxidation was performed to check if carbon deposition was the cause of the
drop in activity. After oxidation, no increase in catalyst activity was observed, confirming the catalyst
does not suffer from carbon deposition.
Compared to the results obtained by P.F. Wolbers, the Ni/Al2O3 shows different behavior in the low
pressure TGA. Wolbers found a weight increase of 160% within 2 hours whereas the low pressure
TGA doesn’t even show a weight increase but a decrease in 5 hours of reaction. Since the same
procedure has been used, these results are not expected.
3.1.3 Determination of minimal fluidization velocity
The Ru/CeZrO2 catalyst will be used in a fluidized bed membrane reactor. An important parameter
in the fluidized bed reactor is the minimal fluidization velocity umf. The theory behind the fluidized
bed concept will be explained later in chapter 4.2.1.
3.1.3.1 Theory on minimal fluidization velocity
The minimal fluidization velocity is the point where the superficial gas velocity is just fluidizing a
bed of particles. A theory exists that a linear correlation is present between the superficial gas velocity
and pressure drop over the bed (Kunii and Levenspiel, 1991). At a certain point, this linear trend
stops and the slope of the pressure drop will decrease, which is defined as the point of minimum
fluidization. This has been visualized in Figure 3.4.
0 5 10 15 20 25 30 35
-4
-2
0
2
4
6
0 5 10 15 20 25 30 35
-4
-2
0
2
4
6
0 5 10 15 20 25 30 35
-4
-2
0
2
4
6
32
Wei
ght d
iffer
ence
(%)
Time (hrs)
CeZrO2
1
3
2
1
Wei
ght d
iffer
ence
(%)
Time (hrs)
Ru/CeZrO2
3
2
1
Wei
ght d
iffer
ence
(%)
Time (hrs)
Ni/Al2O3
0 5 10 15 20 25 30 35
-4
-2
0
2
4
6 CeZrO2 Ru/CeZrO2 Ni/Al2O3
Wei
ght d
iffer
ence
(%)
Time (hrs)
Technische Universiteit Eindhoven University of Technology
19 Novel catalyst for ATR of methane
Figure 3.4 example of Δp versus u0 for uniformly sized sharp sand. Umf is determined as the intersection between the fixed bed line
with the horizontal line W/A (Kunii and Levenspiel, 1991).
For describing the pressure drop in a gas flowing through a packed bed, the Ergun equation is often
used:
2 2
23 3
1 1150 1.75mf mfg mf g mf
mf mf s ps p
U UPL dd
3.1
In which P is the pressure difference [kg/m/s2], L the length of the bed [m], mf the porosity of
the bed [-], g the gas viscosity [Pa·s], s the sphericity of the particle [-], pd the particle diameter
[m], mfU the minimal fluidization velocity [m/s] and g the gas density [kg/m3].
The pressure drop across a bed of particles is given by
1 mf s gP Lg 3.2
In which s is the solid phase density [kg/m3] and g the gravitational constant [9.81 m/s2].
In order to find a correlation for umf, these equations need to be combined a solved.
2 2
23 3
1 11 150 1.75mf mfg mf g mf
mf s gmf mf s ps p
U Ug
dd
3.3
Rearranging the equation by multiplying with
3
2 1g p
mf
d
gives the following rearranged formula:
3 2 2 2
2 3 3 2
1 1.75150g s g p mf p g mf p g mf
g mf g mf g
gd d U d U
3.4
The left side of the equation is also known as the Archimedes number and a part of the right hand
side can be written as the Reynolds number Re d g
, this results in a simpler version of this
equation.
Technische Universiteit Eindhoven University of Technology
20 Novel catalyst for ATR of methane
2
3 3
1 1.75150 Re Remfmf mf
mf mf
Ar
3.5
In many cases, the exact value of mf is unknown. Because of correlations found by several
researchers for different kind of particles, it is possible to predict the minimal fluidization velocity.
When
3
1 mf
mf
is assumed to be constant 1K and 3
1.75
mfassumed to be constant 2K , the minimal
fluidization velocity can be predicted by the following formula:
21 2 1
gmf
g p
u K K Ar Kd
3.6
This correlation will be used later on to compare the obtained results for fluidization with the
literature. Relations for umf are available which are not based on the Ergun equation, but most of the
used correlations in literature are based on this relation.
This correlation gives good predictions for the minimal fluidization velocity at room temperature
and ambient pressure (Kunii and Levenspiel, 1991). Research has been conducted on the effects
temperature and pressure on the minimal fluidization velocity but results are rather inconclusive. A
few conclusions, however, from these studies are:
- εmf increases slightly with a higher operating pressure (1 to 4%)
- umf decreases with a higher operating pressure, this effect is negligible for particles up to
100 µm but becomes significant for particles larger than 300 µm
- εmf increases with temperature for fine particles up to 8%
3.1.3.2 Setup and procedure
The following setup has been used to determine the minimal fluidization velocity.
E-5
Vent
N2
PI
Oven
3 l/min
TIC
Figure 3.5 Schematic of the minimal fluidization setup
The reactor had an internal diameter of 2.5 cm. A pressure indicator was connected close the bottom
of the bed and a mass flow controller was installed for nitrogen with a flowrate up to 3 Nl/min. The
temperature was recorded with a thermocouple inside the oven.
Technische Universiteit Eindhoven University of Technology
21 Novel catalyst for ATR of methane
To determine the minimal fluidization velocity, the gas velocity was increased stepwise and the
pressure drop was recorded for each step. The minimal fluidization was determined as the
intersection between the fixed bed line with the horizontal line W/At, because of the size distribution
of the used particles, Figure 3.4 (Kunii and Levenspiel, 1991).
Figure 3.6 Δp versus u0 for the case of CeZrO2 particles with particle size 90-125 µm at 20 °C (l). Because of the wide distribution
of particle size, the minimal fluidization velocity is determined as the intersection between the fixed bed line with the horizontal line
W/At (r).
Five fractions of particle sizes were tested ranging from 90 to 355 µm and the minimal fluidization
velocity has been determined at several temperatures in a range from 20 to 600 °C.
3.1.3.3 Results
The determined minimal fluidization velocities can be found in Table 3.1 and is visualized in Figure
3.7 and Figure 3.8.
Table 3.1 minimal fluidization velocity for several particle size distributions of CeZrO2
Particle size distribution (µm)
90-125 125-180 180-250 250-315 315-355
T (°C)
umf (cm/s)
T (°C)
umf (cm/s)
T (°C)
umf (cm/s)
T (°C)
umf (cm/s)
T (°C)
umf (cm/s)
20 1.73 20 3.31 20 6.53 20 -1 20 -1
109 1.58 103 2.68 105 5.54 105 9.86 104 -1
214 1.40 214 2.71 200 4.38 210 8.32 208 11.37
311 1.26 309 2.15 304 3.83 324 7.39 308 10.41
390 1.21 393 2.10 387 3.83 387 6.37 386 9.46
484 1.16 484 2.03 488 3.70 488 5.66 486 7.77
585 1.17 574 1.93 584 3.26 583 5.06 589 7.16 1 no fluidization below 3 Nml/min N2 (range of the MFC)
0.00
2.00
4.00
6.00
8.00
10.00
12.00
14.00
16.00
0 1 2 3 4 5
pre
ssu
re d
rop
Δp
(mb
ar)
gas velocity u0 (cm/s)
Forward
Backward
y = 0.2432x + 14.085R² = 0.9734
y = 8.0731x + 0.5391R² = 0.9978
0.00
2.00
4.00
6.00
8.00
10.00
12.00
14.00
16.00
18.00
0 1 2 3 4 5
pre
ssu
re d
rop
Δp
(m
bar
)
gas velocity u0 (cm/s)
Technische Universiteit Eindhoven University of Technology
22 Novel catalyst for ATR of methane
Figure 3.7 The determined minimal fluidization velocity versus the temperature for several fractions of CeZrO2.
Figure 3.8 The determined minimal fluidization velocity versus the average particle size of the fractions of CeZrO2.
The results of the determination of the minimal fluidization velocity show expected behavior. From
literature it is known that the minimal fluidization velocity decreases with increasing temperature
and with decreasing particle size.
3.1.3.4 Comparison with literature
The obtained results have been compared to results published in literature (Kunii and Levenspiel,
1991). Several investigators have reported correlations between the gas and particle properties with
the minimal fluidization velocity. The most used form of predicting umf is by using equation 3.6.
The error in these correlations is about 40%. The correlations to compare the results with are
reported to be correlations for Geldart B particles, which is also the classification for the particles
used in this work.
0 100 200 300 400 500 6000
2
4
6
8
10
12Particle size (m)
90-125 125-180 180-250 250-315 315-355
min
imal
flui
diza
tion
velo
city
um
f (cm
/s)
Temperature (deg C)
100 150 200 250 300 3500
2
4
6
8
10
12
min
imal
flui
diza
tion
velo
city
um
f (cm
/s)
Average particle size (m)
Temperature (C) 20 100 200 300 400 500 600
Technische Universiteit Eindhoven University of Technology
23 Novel catalyst for ATR of methane
The acquired data has been compared to the following investigators values:
Table 3.2 Values of several investigators of the two constants in equation (Kunii and Levenspiel, 1991)
Investigator K1 K2
Thonglimp (1981) 31.6 0.042
Richardson (1971) 25.7 0.0365
Wen and Yu (1966) 33.7 0.0408
Grace (1982) 27.2 0.0408
For the model prediction the temperature effects for the viscosity and density of the gas are taken
into account. Sutherland’s formula has been used to correct for gas viscosity and for the gas density
the ideal gas law was applied.
3/2
00
0
T C TT C T
3.7
Where is the dynamic viscosity at temperature T (Pa·s), 0 is the dynamic viscosity at reference
temperature T0 (Pa·s), T the input temperature (K), 0T the reference temperature (K) and CSutherland’s constant (-). This correlation is proven to be valid for temperatures between 0 < T < 555
K with an error below 10%.
Applying these models and plotting the data gives the following results:
Figure 3.9 Experimental value for umf in comparison with theoretical models for different particle size distributions
100 200 300 400 500 600
1.0
1.5
2.0
2.5
100 200 300 400 500 600
2
3
4
5
100 200 300 400 500 600
3
4
5
6
7
8
9
100 200 300 400 500 600
56789
101112131415
100 200 300 400 500 6006789
1011121314151617181920
min
imal
flui
diza
tion
velo
city
um
f (cm
/s)
Temperature (C)
90-125 m
125-180 m
min
imal
flui
diza
tion
velo
city
um
f (cm
/s)
Temperature (C)
315-355 m
250-315 m180-250 m
min
imal
flui
diza
tion
velo
city
um
f (cm
/s)
Temperature (C)
m
inim
al fl
uidi
zatio
n ve
loci
ty u
mf (
cm/s
)
Temperature (C)
Thonglimp Richardson Wen and Yu Grace Experimental
min
imal
flui
diza
tion
velo
city
um
f (cm
/s)
Temperature (C)
Technische Universiteit Eindhoven University of Technology
24 Novel catalyst for ATR of methane
To calculate the error of the models, parity plots were made which can be found in the Appendix.
With these parity plots, the coefficient of determination (R-squared) was calculated in which the
observed value vs. the predicted model value was fitted with the line y=x. This gives an indication of
how well the correlation predicts the actual experimental minimal fluidization velocity. The closer
the value is to 1, the better the fit is, thus the better the correlation is predicting.
Table 3.3 Coefficient of determination (R-squared) for the observed data and predicted model values versus the fitted line y=x
Particle fraction
(µm)
Thonglimp (1981) Richardson (1971) Wen and Yu
(1966)
Grace (1982)
90-125 0.812 0.800 0.562 0.699
125-180 0.770 0.550 0.900 0.310
180-250 0.631 0.320 0.916 -0.003
250-315 0.451 -0.080 0.880 -0.638
315-355 -0.311 -1.467 0.667 -2.655
It can be concluded, that the correlation of Wen and Yu gives the best prediction for the fluidization
behavior for all particle fractions except for the smallest fraction. A reason for this deviation is most
probably that the correlations do not take into account the temperature dependency of the bed
voidage at minimal fluidization. For the experimental work, the Wen and Yu correlation has been
used to predict the minimal fluidization velocity. The particles which are used in the experiments
are of the size 125-250 µm so no large deviations are expected.
3.1.4 Fluidization behavior
The catalyst will be used with a filler in the FBMR to extend and dilute the bed. Therefore it is
important to know the fluidization behavior of the catalyst with the filler. The following mixture has
been used as filler as this composition was predicted to have the same fluidization behavior:
Table 3.4 Composition of the filler used for the FBMR, developed by Hybrid Catalysis
Zirconia based filler
250 sieve fraction
Component Weight%
ZrO2 60-70
SiO2 28-33
Al2O3 <10
Segregation of a binary mixture of particles occurs when there is a substantial difference between
their drag force per unit weight. A quantitative measure for segregation was given by Tanimoto et
al. (Tanimoto et al., 1981) and was later modified by Hoffman et al. (Hoffmann and Romp, 1991):
1/3
,
,
0.8 0.8p j j
p f f
dY
d
3.8
Where Y is the segregation distance, pd the particle diameter, the particle density and j stands
for jetsam particles (sinking particles) and f for flotsam particles (rising particles).
If the segregation distance is close to zero, no segregation effects occur. The filler is designed to have
similar properties to the catalyst. For this given system, the particle diameter of both particles is
assumed to be the same, both are the 125-250 µm sieve fraction and the particle density of the filler
Technische Universiteit Eindhoven University of Technology
25 Novel catalyst for ATR of methane
is adjusted to be the same as the catalyst (2100-2300 kg/m3), hence the mixture. From theory a
segregation distance near zero is obtained, thus no segregation effects are expected.
To evaluate the fluidization behavior with this filler, the support (CeZrO2) was mixed with the filler
in a 30 to 70 ratio (as will be used in the FBMR) and fluidized for a day in a 2D fluidized bed. Pictures
have been taken every hour to ensure that the particles were still mixed properly and no segregation
effects occur.
Table 3.5 Pictures of the segregation test of CeZrO2 (yellow) with the zirconia based filler particles (white) in a 2D fluidized bed
At start of fluidization After 8 hours After 25 hours
From these pictures it was concluded that in general no segregation effects occur under fluidization
at room temperature. The particles were well mixed and stayed well mixed over time. A cluster of
filler particles was detected in the bottom, which is most probably caused by the particle size
distribution. It could be that there is a higher fraction of smaller particles in the filler, it could also
be that the particle density is less uniform, thus having a fraction with a higher particle density which
prefer staying at the bottom. This amount is however that minimal that it is expected it has no
significant effect on the mixing of the particles in the FBMR.
Technische Universiteit Eindhoven University of Technology
26 Novel catalyst for ATR of methane
3.2 Catalyst stability
One important parameter for the catalyst, designed for ATR, is the stable activity (methane
conversion) over long time. The Rh/ZrO2 catalyst has already been investigated by P. Wolbers and a
reproducibility experiment has been conducted on this catalyst.
A long-term (>100 hours) stability test has also been performed for the new Ru/CeZrO2 catalyst to
evaluate the stability.
3.2.1 Experimental
The following setup has been used to determine the catalyst stability and activity.
Figure 3.10 Schematic of the experimental setup to determine the catalyst stability in a packed bed reactor of quartz
The reactor that has been used is a quartz u-shaped reactor with an internal diameter of 6 mm. A
narrowing was made in the tube to provide for a possibility to create a packed bed. Glasswool was
used in the narrowing and a packed bed of approximately 30 mg of catalyst mixed with quartz
particles (0.3<dp<0.5 mm) in a ratio of 1:7 to dilute the bed and to extend the active bed height, was
created. The reactor was then filled with quartz particles (1<dp<1.7 mm) to decrease the gas fraction
inside the reactor and to increase the gas velocity, with the aim to prevent gas phase reactions as can
be seen in Figure 3.11. The u-shaped reactor was placed in a fluidized bed sand oven.
Figure 3.11 Schematic of the quartz u-shaped reactor at the height of the catalyst and sand oven (l) actual reactor filled with diluted
Ru/CeZrO2 (bed height = ±1 cm) (r)
N2
FIC102
CH4
FIC202
CO FIC302
H2
FIC402
FV101
FV201
FV301
FV401
Steam
CoolerHEX-807
O2
FIC702
FV701
CH4
FIC802
FV801
N2
FIC902
FV901
CO2
FV303
FV501
FV600
TC501
TC502
bypass
TI609
TIC610
TI611
FI601
FI607
FV608
Exhaust
GCTI
612
TI613
TI603B
TIRC603A
PRV 601
Exhaust
PS601
PS601
TI614
PS604
TI605
Thermocouples Catalyst Fluidized
Bed
Technische Universiteit Eindhoven University of Technology
27 Novel catalyst for ATR of methane
The mass flow controllers have been calibrated prior to the experiments. The steam was fed using a
HPLC pump followed by an evaporation unit. The lines of the steam were provided with tracing to
prevent steam condensation. The outlet gas stream was analyzed with a Galaxie microGC. The
microGC has been calibrated to measure in the expected range. Water traps were installed to prevent
steam/water entering the GC.
To have comparable results, both catalysts have been subjected to identical conditions. The same
procedure has been followed for all experiments.
1) The reactor was heated up to 650 °C with a ramp of 2 °C/min
2) Activation/reduction of the catalyst by feeding 20% H2 / 80% N2 (0.0134 mol/min)
3) Autothermal reforming of methane using 16% CH4, 30% steam, 7% O2 and 47% N2 (0.07
mol/min). This results in an SCR of 1.9, OCR of 0.44 and an NCR of 3.
3.2.2 Limitations of the system
To determine the catalyst performance, it is required that the system is operated in the kinetically
controlled regime. This means the system will not be able to reach thermodynamic equilibrium
conversion due to the low residence time of the reactants. By applying the kinetically controlled
regime it is possible to evaluate the catalytic performance which is then limited by the reaction rate
and not by other limitations.
To confirm the reaction rate is indeed dominant, the external mass transfer contribution should be
evaluated. The Mears criterion (Mears, 1971), Eq. 3.9, indicates whether the (methane) conversion is
unaffected by external diffusion limitations.
4
4
exp* 0.15CH
CH c c c
r kc k a k
3.9
In this equation, 4CHr is the methane reaction rate [mol/(m3s)],
4CHc is the methane concentration
[mol/m3], ck is the mass transfer coefficient [m/s], ca is the external surface-to-volume ratio of the
particle [m2/m3], expk is the experimental kinetic rate constant [1/s] and *ck is the intrinsic mass
transfer coefficient [1/s].
The reaction rate, combined with the methane concentration is captured in expk and is determined
by experimental results.
4 44
4 4
exp
y (1 ) /
/CH tot CH cat catCH
CH CH
mrk
c y p RT
3.10
In this equation, 4
yCH is the molar fraction of methane in the feed [-], tot is the total molar feed
flow [mol/s], 4CH the methane conversion [-], cat the catalyst density [kg/m3], the bed porosity [-
Technische Universiteit Eindhoven University of Technology
28 Novel catalyst for ATR of methane
], catm the catalyst mass [kg], p the feed pressure [Pa], R the gas constant [J/mol/K] and T the gas
temperature in the bed [K].
The intrinsic mass transfer coefficient *ck [1/s] is given by:
* vc c c c
p
Dk k a Sh ad
3.11
With vD the binary diffusivity [m2/s], pd the particle diameter [m], Sh the Sherwood number,
which for a packed bed reactor is described by the following expression (Marra et al., 2013):
1/3 1/22 1.5 (1 )ReSh Sc 3.12
Where Sc is the Schmidt number and Re the Reynolds number.
ScD
3.13
Re(1 )vL
3.14
In which is the dynamic viscosity of the gas [Pa·s], the gas density [kg/m3], D the binary mass
diffusivity [m2/s], v the superficial gas velocity [m/s] and L the height of the bed [m].
With the used settings the Mears criterion has a value of 0.57 which is higher than the 0.15 and
therefore the external diffusion limitation is not negligible. The setup did not allow for operation
free of external mass transfer limitations due to the high pressure drop. External mass transfer
limitations should be taken into account, interpreting the experimental data.
The parameters used for the experiment as described in the experimental chapter can be found in
the Appendix.
The binary mass diffusivity has been calculated by the correlation of Fuller et al. (Poling et al., 2001).
The dynamic viscosity has been calculated for the feed mixture at 650 °C using the correlation of
Sutherland.
To evaluate internal mass transfer limitations, the effectiveness factor is used which is given by the
following equation (Marra et al., 2013):
1 1 1tanh 3 3
3.15
The effectiveness factor uses the Thiele modulus which compares the diffusion time versus the
reaction time:
exp
6pd k
D 3.16
Technische Universiteit Eindhoven University of Technology
29 Novel catalyst for ATR of methane
Where pd is the catalyst particle diameter [m], expk the experimental kinetic rate constant [1/s] and
D the binary mass diffusivity [m2/s].
The effectiveness factor found for this system is 0.98 which means that the internal mass transfer
limitations can be neglected.
3.2.3 Results
3.2.3.1 Reproducibility study on Rh/ZrO2
The Rh/ZrO2 catalyst has been investigated by P. Wolbers. An activity experiment has been
conducted to reproduce a part of the results. The procedure as mentioned before has been followed
and the results can be found in Figure 3.12 together with the results of P. Wolbers.
Figure 3.12 Methane conversion as a result of ATR of methane using a Rh/ZrO2
It is quite clear that the conversion obtained in this work is lower than the values found by P.
Wolbers. Both experiments followed the same procedure and the weights of the catalyst samples
were 29.4 mg for Wolbers and 31.9 mg for this work. However, the mass flow controllers were not
calibrated correctly, resulting in a higher flow than 0.07 mol/min, which had been detected after the
experiments. Another difference that is observed was the temperature increase during reaction. P.
Wolbers reported a temperature increases of 150 °C, resulting in an effective temperature >800 °C,
whereas during the stability test in this work a temperature increase of only 50 °C could be detected.
The placement of the thermocouple is important in monitoring the temperature, it could be that the
thermocouple in the experiment of P. Wolbers were closer to the bed, giving a more accurate
description of the temperature. A temperature increase has a positive effect on the thermodynamic
equilibrium of the system, which is the most probable explanation for the difference in methane
conversion that is observed. This is also powered by the fact that in the beginning (t < 10 min) the
methane conversion is almost similar.
3.2.3.2 Stability test on Ru/CeZrO2
The goal of the new Ru/CeZrO2 catalyst is to operate ATR of methane in a FBMR. It is therefore
important to confirm if the catalyst is suitable for this operation. The stability test has been
performed using the method as described in chapter 3.2.1 and it has been conducted on the support
and the supported catalyst particles.
0 50 100 150 200 250 300 3500
20
40
60
80
100
Met
hane
con
vers
ion
(%)
Time (min)
P. Wolbers this work
Technische Universiteit Eindhoven University of Technology
30 Novel catalyst for ATR of methane
Support
The results of ATR of methane with the support particles can be found in Figure 3.13. From this data
it can be concluded that the support does not influence the reaction. One could argue, that there is
a small production of CO2, but this could be the error in the measurements of the GC or gas phase
reaction between methane and oxygen. Most probably it is the error of the measurement as the
methane did not show any decrease.
Figure 3.13 Outlet composition on the stability test on CeZrO2
Because the support did not show any apparent reaction, it was decided no test will be conducted on
an empty reactor as it is expected to show similar results.
Ru/CeZrO2
After it was confirmed the support does not initiate the ATR of methane, the catalyst has been tested
to check if the designed catalyst serves its purpose. The same procedure has been followed and the
results are shown in Figure 3.14.
Figure 3.14 Outlet composition (l) and methane conversion (r) on the stability test on Ru/CeZrO2
0 20 40 60 80
0
10
20
30
40
50
60
Methane Oxygen Hydrogen Carbon dioxide Nitrogen blank
Out
let c
ompo
sitio
n (v
ol%
)
Time (min)
T = 650 °C
0 200 400 600 800 1000 1200 1400
0
10
20
30
40
50
60
70
Out
let c
ompo
sitio
n (v
ol%
)
Time (min)
Methane Oxygen Hydrogen Carbon dioxide Carbon monoxide Nitrogen
0 200 400 600 800 1000 1200 14000
5
10
15
20
25
Met
hane
con
vers
ion
(%)
Time (min)
Technische Universiteit Eindhoven University of Technology
31 Novel catalyst for ATR of methane
It can be noticed right away is that the activity of the catalyst is not stable and that within one day,
the activity drops to almost zero. Furthermore, the methane conversion at maximum is around 23%
which is not close to the expected >50% achieved by the Rh/ZrO2 catalyst. It should be mentioned
that the catalyst loading is different: 1.4wt% Rh on ZrO2 and 2wt% Ru on CeZrO2. As the sample
weights were the same (approximately 30 mg), this means the difference between the active
substances is already 30%, but in contrast to the results, the Ruthenium catalyst with a higher
loading shows lower conversion.
The density of the new Ruthenium catalyst is 4400 kg/m3 and is 1000 kg/m3 higher than the
Rhodium catalyst. Because the particles used for both test were in the same range (125-250 µm),
which implies that the effective surface area of the Ruthenium-based catalyst is about 20% lower
than the rhodium catalyst. BET analysis on both catalyst was conducted to confirm if this is the case.
The results of the BET analysis can be found in Table 3.6.
Table 3.6 BET analysis on the rhodium and ruthenium based catalyst and supports
Sample Pore Volume (cm3·g-1)
BET Surface Area (m2·g-1)
Calcined ZrO2 0.2419 64
1.4 wt% Rh/ZrO2 0.2188 57
1.4 wt% Rh/ZrO2 after the experiments 0.1546 37
Ce0.75Zr0.25O2 0.1092 103
2 wt% Ru/CeZrO2 0.1027 88
From this data it can be concluded that ruthenium based catalyst has a higher surface area. The pore
volume for the ruthenium-based catalyst is significantly lower than for the rhodium-based catalyst.
This could indicate that there are less active sites available for the ruthenium catalyst, which could
contribute to the lower activity observed.
Several articles on the steam methane reforming on ruthenium and rhodium based catalysts have
been published. It had been reported that rhodium is a far more active catalyst for steam methane
reforming and that ruthenium benefits greatly from high metal loadings (Kusakabe et al., 2004;
Liguras et al., 2003).
The stability test was repeated and similar results were obtained. Another parameter which could be
causing deactivation of the catalyst is the temperature. The temperature is monitored by three
thermocouples as can be seen in Figure 3.11. The temperature profiles of the thermocouples before
and after the packed bed are added as Figure 3.15.
Technische Universiteit Eindhoven University of Technology
32 Novel catalyst for ATR of methane
Figure 3.15 Temperature profile over time in the second batch stability test on Ru/CeZrO2
In the beginning of the experiment, the temperature seems to be fluctuating a lot with a ΔT of about
30 °C, when the catalyst is still active. Although the temperature is stable in the middle of the
experiment, the methane conversion is still decreasing. The decrease in activity can therefore not be
caused by a temperature instability. It is concluded that the catalyst is not suitable for ATR in a
packed bed due to a decrease in activity.
The stability test on the activity of the catalyst was also done by Hybrid Catalysis. Their results
showed a stable activity and the graph is added as Figure 3.16.
Figure 3.16 Stable hydrogen production as found by Hybrid catalysis with ATR on the Ru/CeZrO2 catalyst (green line)
In comparison to the trend observed in the tests in this work, the hydrogen production is stable. The
main difference between the tests are the flows used for the experiments. As this experiment is
important for the catalyst choice in the fluidized bed, the test has been redone at the same conditions
and the same weight hourly space velocity (WHSV).
Mass flowWHSV=
Catalyst mass 3.17
The WHSV that was asked to be evaluated was 22 h-1, this is inspired on the pilot plant of Hygear.
To give a reference, the calculated WHSV used for the first stability tests were around 500 h-1. Next
0 200 400 600 800 1000 1200
630
640
650
660
670
680
690
Tem
pera
ture
(°C
)
Time (min)
Before the bed After the bed
Technische Universiteit Eindhoven University of Technology
33 Novel catalyst for ATR of methane
to the inlet flow, the inlet composition was also changed. Instead of a Nitrogen to Carbon Ration
(NCR) of 3, an NCR of 1.66 was proposed, meaning the feed was less diluted.
Figure 3.17 Outlet composition (l) and methane conversion (r) on the stability test on Ru/CeZrO2 with a WHSV of 22 h-1
The results in Figure 3.17 show a clear decrease in methane conversion over time. The catalyst
deactivation has a smaller slope due to the lower flows used. An increase of activity has been observed
when the catalyst had been subjected to oxidation to rule out carbon deposition. The increase in
activity is small, indicating deactivation is not due to carbon deposition.
Since the catalyst will be used in a fluidized bed and deactivation most probably occurs due to poor
heat transfer in the packed bed, it was decided to continue using the Ruthenium catalyst as the heat
transfer is much better. Ruthenium is about 20 times cheaper than rhodium and is worthwhile
testing for the fluidized bed membrane reactor concept as this is a preliminary test to check if the
concept is viable for ATR.
No kinetics study has been performed on the catalyst. For concept viability of the FBMR reactor, the
catalyst can be used as the WHSV will be much lower (in the range of 0.0665 h-1) and a lot more
catalyst is present in the system and no decrease in activity is expected to be observed due to the
excess of catalyst.
0 500 1000 1500 2000 2500 3000 3500
0
20
40
Out
let c
ompo
sitio
n (v
ol%
)
Time (min)
Oxygen Nitrogen Hydrogen Methane Carbon monoxide Carbon dioxide
0 500 1000 1500 2000 2500 3000 35000
20
40
60
Met
hane
con
vers
ion
(%)
Time (min)
oxidation
Technische Universiteit Eindhoven University of Technology
34 Novel catalyst for ATR of methane
3.3 Conclusion
It can be concluded that the Ru/CeZrO2 is mechanically stable. Fluidization tests have been
conducted at room temperature and high temperature and no change in particle size distribution
has been observed. From the thermo gravimetric analysis at the conditions of steam methane
reforming no carbon deposition has been detected. This is confirmed in the stability test of
Ru/CeZrO2 where oxidation did not improve catalyst activity.
The minimal fluidization has been determined for the CeZrO2 particles and the correlation of Wen
and Yu is describing the minimal fluidization velocity for these particles the best. These CeZrO2
have also been fluidized with a zirconia based filler and after 25 hours, no segregation was observed.
The stability tests in the packed bed reactor showed that the activity of the catalyst in terms of
methane conversion is not stable. Similar results were obtained with a second stability test with a
fresh batch of catalyst. With a lower flow, the catalyst also showed deactivation. Although the catalyst
activity is decreasing over reaction time, it will be used later on to test the fluidized bed membrane
reactor for ATR of methane. Excess of catalyst will be used and flows will be much lower so no
decrease in activity observed is expected.
Next to the ruthenium based catalyst, the rhodium based catalyst has been tested to see if similar
results are obtained as P. Wolbers. Difference in methane conversion were observed which can be
explained by incorrect feed flows and a temperature difference. Taking this into account, the
methane conversion is relatively close to the conversion found by P. Wolbers. The kinetic study
performed by P. Wolbers will therefore be used later in this work to compare reactor concepts
theoretically.
Technische Universiteit Eindhoven University of Technology
35 Evaluation of reactor concepts for ATR
4. Evaluation of reactor concepts for ATR
The membranes have proven to be selective to hydrogen and the catalyst have shown to enhance the
autothermal reforming of methane. Although the Ru/CeZrO2 catalyst seems to have stability issues,
the deactivation is most probably caused by hot spots within the bed. As this catalyst will be used for
the fluidized bed with high internal heat exchange, no problems are expected with the activity of the
catalyst. From an economical point of view, the rhodium catalyst is approximately 20 times more
expensive than the ruthenium based catalyst. It is worthwhile to carry on testing the ruthenium
based catalyst.
So for the FBMR, the ruthenium based catalyst was used as a large volume of particles was required
(0.5 kg of catalyst mixed with 3 kg of filler material). In this case it is not expected that catalyst
degradation will be seen due to the excess of catalyst compared to the inlet flows used.
Technische Universiteit Eindhoven University of Technology
36 Evaluation of reactor concepts for ATR
4.1 Microreactor
The first reactor concept which was evaluated is the microreactor. For the microreactor different wall
coatings were evaluated, one developed by Tecnalia and one developed by T. Janssen at TU/e. Both
used the rhodium based catalyst with different wall thickness and metal concentration. The Tecnalia
microreactor was coated with 4.2 mg of 20 wt% Rh/ZrO2 with a wall thickness of 5 µm whereas the
microreactor coated by T. Janssen had a 16.8 mg 1.4 wt% Rh/ZrO2 coating and a wall thickness of
235 µm.
Another microreactor concept that was evaluated was the packed bed membrane microreactor
(PBMMR).
4.1.1 Theory
The microreactor concept is a promising reactor type due to its superior external heat transfer, the
ability to capture intermediate products and the simplicity in scaling up (multiplying the number of
channels).
Microreactors are characterized by their high surface area-to-volume ratios in their microstructured
regions that contain tubes or channels. This high area-to-volume ratio reduces heat and mass
transfer resistances often found in larger reactors. Consequently, surface-catalyzed reactions can be
greatly facilitated, hot spots in exothermic reactions can be eliminated and in many cases, highly
exothermic reactions can even be operated isothermally. Microreactors are therefore often used to
study the intrinsic kinetics of reaction.
Another advantage of microreactors is their safety when producing toxic or explosive intermediates.
A leak or explosion in a single channel will do minimal damage because of the small quantities of
material involved. Next to that, the microchannels provide for shorter residence times and narrower
residence time distributions allowing for better operational control.
A microreactor with microchannels can be considered to behave like either a plug flow reactor or
laminar flow.
4.1.2 Preparation and procedure
Both tests were carried out in a microreactor consisting of six channels of the dimensions 1 x 1 x 13
mm. A schematic of the microreactor is added as Figure 4.1. In the case of the wall coated
microreactors, no membrane was present and the permeate side was flat. So the retentate side was
only sealed. In the case of the PBMMR, the schematic is representative.
Technische Universiteit Eindhoven University of Technology
37 Evaluation of reactor concepts for ATR
Figure 4.1 Schematic of a microreactor (A. L. Mejdell et al., 2009a)
4.1.2.1 Wall coated microreactor
Tecnalia coated microreactor
The wall coated microreactor of Tecnalia consisted of 4.2 mg 20 wt% Rh/ZrO2 with a wall thickness
of 5 µm which was deposited via traditional impregnation method.
T. Janssen coated microreactor
The wall coating T. Janssen used was prepared by co-deposition method using a ZrO2/Al2O3 sol. 16.8
mg of 1.4 wt% Rh/ZrO2 was deposited in the microreactor channels which was then calcined in the
oven at 650 °C for 4 hours. This resulted in a wall thickness of 235 µm. It has to be noted that the
wall thickness was not uniform.
4.1.2.2 Packed bed membrane microreactor
To create a packed bed, glasswool was inserted just before and after the channels on the retentate
side preventing any particles from moving. 45 mg of Ru/CeZrO2 was inserted in the channels to
create the packed bed without overflow out of the channels. The membrane was sealed on top of it
and the permeate was kept free.
The reaction was performed using the kinetic setup (Figure 3.10). The procedure that has been
followed is the same as mentioned in chapter 3.2.1. The only difference is the total flow rate was set
at 0.02 mol/min due to limitations of the setup and the microreactor in terms of pressure drop.
4.1.3 Results
4.1.3.1 Rh/ZrO2 wall coated microreactor
Tecnalia coated microreactor
The wall coated microreactor by Tecnalia showed no hydrogen production. A low methane
conversion was observed due to oxidation of methane. Several settings were changed to improve the
reaction. The total flow rate was decreased to increase the residence time but this didn’t improve the
conversion. The temperature was also increased to shift the thermodynamic equilibrium of the SMR
reaction, but no change was observed.
Technische Universiteit Eindhoven University of Technology
38 Evaluation of reactor concepts for ATR
T. Janssen coated microreactor
Similar to the results of the microreactor of Tecnalia, no hydrogen production was observed. There
was some conversion of methane but also in this case, this is probably due to oxidation. Decrease of
the flowrate or increase of temperature did not improve the performance.
Due to the low conversions observed, it was decided to perform the membrane assisted microreactor
tests with a packed bed inside the microchannels instead of wall coating to increase the apparent
reaction rate.
4.1.3.2 Rh/ZrO2 Packed bed membrane microreactor
During the test, the outlet stream was continuously analyzed by a SICK gas analyzer. The results of
the test are presented in Table 4.1 and Figure 4.2.
Table 4.1 Summary of the ATR test in a packed bed membrane microreactor with a SCR of 1.9, OCR of 0.44 and NCR of 3 at
different temperatures and pressures in- and excluding separation with vacuum on the permeate side
Experiment R R+S R R+S R R+S R R+S R R+S
Temperature [°C] 550 550 600 600 600
Reactor pressure [bar] 1.0 1.5 1.0 1.4 1.9
Methane conversion [%] 28 28 28.5 29.9 27.4 26.9 33.6 33.1 27.7 28.6
Oxygen conversion [%] 100 100 100 100 100 100 100 100 100 100
Thermodynamic
equilibrium
methane conversion
[%] 52.1 39.6 78.2 69.1 59.9
Also for this case the Mears criterion and effectiveness factor have been evaluated. The microreactor
has been assumed to behave like a packed bed, the Sherwood correlation for the packed bed has
therefore been used (Equation 3.12). The Mears criterion had a value of 0.53 which is above 0.15 so
external mass transfer limitations are present, which could not be avoided, due to limitations to the
setup (high pressure drop. The effectiveness factor had a value of 0.98 which is close to 1, so it’s
assumed the system is not suffering from internal mass transfer limitations.
Figure 4.2 outlet composition of the ATR in a PBMMR with varying temperatures and pressures
During the test there were quite some problems maintaining the pressure drop within limits causing
for fluctuations in the results in Figure 4.2. Next to the experimental methane conversion, the
methane conversion based on thermodynamic equilibrium is given. A clear difference is observed
between the actual and the expected methane conversion. What is also quite remarkable is the low
hydrogen production. The maximum value of hydrogen in the outlet stream is about 5% which is
0 50 100 150 200 250 3000
5
10
15
20
25
R1.4 bar600 °C
R&S1.0 bar600 °C
R1.0 bar600 °C
R1.5 bar550 °C
R&S1.5 bar550 °C reaction and
separation1.0 bar550 °C
Out
let c
ompo
sitio
n (v
ol%
)
Reaction time (min)
CO CO2 CH4 O2 H2
reaction1.0 bar550 °C
R&S1.4 bar600 °C
R&S1.9 bar600 °C
R1.9 bar600 °C
Technische Universiteit Eindhoven University of Technology
39 Evaluation of reactor concepts for ATR
extremely low. The oxygen however, seems to fully react with either the methane in the system or
the hydrogen. So a possible explanation for the low hydrogen production could be the burning due
to oxygen.
The results are not as expected, whether it’s the inlet flow of steam, the pressure in the system or the
temperature in the system, no distinct deviation could be found to the desired settings of the system.
Also in the preparation no deviations were observed, the catalyst was trapped within the physical
constraints and the membrane was placed on top of it. No sealing errors were found and thus it’s
impossible to draw a clear conclusion to why the hydrogen production and methane conversion are
unexpectedly low.
Because of the low hydrogen production, the partial pressure of hydrogen is very low and no
hydrogen permeation through the membrane was observed. The driving force, the transmembrane
hydrogen partial pressure, is almost zero.
4.1.4 Conclusion
From experimental data it seems that a microreactor is not suitable for autothermal reforming. In
all cases the hydrogen production is very low or even not noticeable. Reasonable conversions were
expected as mass transfer limitations should not play a big role.
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40 Evaluation of reactor concepts for ATR
4.2 Fluidized bed membrane reactor
The second reactor concept that has been evaluated is the fluidized bed membrane reactor. A typical
fluidized bed membrane reactor consists of a catalytic bed in the bubbling or turbulent regime in
which a bundle of hydrogen-selective membranes are submerged.
For this experiment the fluidized bed consisted of 0.5 kg of 2wt% Ru/CeZrO2 and 3.5 kg of zirconia
based filler (for more information see chapter 3.1.4). This was accompanied by five tubular Pd-
Ag/ZrO2 membranes developed by Tecnalia. A test was also performed using commercial hydrogen
selective membranes produced by REB to evaluate the catalyst performance in the FBMR.
Next to autothermal reforming, the stability of the membranes within a bed of fluidizing particles is
evaluated. This is important for the selectivity of the membrane and therefore the purity of the
extracted hydrogen.
4.2.1 Theory
The fluidized bed concept is a bed of small particles suspended and kept in motion by an upward
flow of fluid. Therefore, the fluidized bed has excellent heat and mass transfer within the bed.
4.2.1.1 Fluidization Regimes
When a bed of particles is exposed to a flow of gas, a few particles will vibrate but still within the
same height as the rest of the bed at rest. This is also called the fixed bed (Figure 4.3,A). With
increasing velocity, a point is reached where the drag force equals the gravitational force and the
voidage of the bed increases slightly. This is the point of minimal fluidization with a corresponding
minimal fluidization velocity umf (B). Increasing the gas flow further, bubbles start to form, which is
called a bubbling bed (C). Increasing the gas velocity further will cause bubbles to coalescence, thus
creating very large bubbles. It could happen that the diameter of the bubble will exceed the diameter
of the reactor and slugs will form, this is the slugging bed (D). If the gas velocity exceeds the terminal
velocity and the particles are unable to fall back, a turbulent bed is observed where turbulent motion
of solid clusters and voids of gas of various sizes and shapes are observed (E). If the gas velocity is
increased even further, the bed will become an entrained bed where the forces of the gas flow are
that high that the particles are blown out of the reactor (F).
Figure 4.3 Schematic representation of fluidized beds in different regimes (Kunii and Levenspiel, 1991)
Technische Universiteit Eindhoven University of Technology
41 Evaluation of reactor concepts for ATR
4.2.1.2 Geldart’s classification of particles
Next to the gas velocity, the particles itself influence the fluidization behavior of the bed. For this,
Geldart defined four main types of particles which are mainly distinguished by particle size and
density (Baker and Geldart, 1978; Geldart and Abrahamsen, 1978; Geldart, 1973). These four types
of particles have their own characteristics and are categorized as follows and are visualized in Figure
4.4:
Geldart A particles
Group A particles are also depicted as aeratable particles. Particle characteristics are small mean
particle size (dp < 30 µm) and/or a low particle density (ρp < 1.4 g/cm3). This results in easy fluidizable
particles with smooth fluidization at low gas velocities without the formation of bubbles. The Geldart
A particles are able to form bubbles in the bed, the minimal velocity at which bubbles form (umb) is
always higher than the minimal fluidization velocity (umf).
Geldart B particles
Group B particles are depicted as sandlike particles and sometimes bubbly particles. The particle
size ranges from 150 < dp < 500 µm and densities between 1.4 < ρp < 4 g/cm3. In contrast to Geldart
A particles, when fluidization is reached, the excess gas will be transported in the form of bubbles.
This means umb is close to umf.
Geldart C particles
Group C particles are cohesive particles, or very fine powders. The size of the particles is usually less
than 30 µm and they are difficult to be fluidized. This is because the interparticle forces are relatively
large compared to the forces resulting from the gas. In small diameter beds, group C particles will
give rise to channeling.
Geldart D particles
Geldart D particles are spoutable particles and usually are very large or very dense particles. They are
rather difficult to fluidize and as velocity increases, a jet can be formed in the bed causing particles
to be blown out with the jet in a spouting motion. Therefore it’s key to have an even gas distribution,
to avoid spouting behavior and channeling.
Figure 4.4 Diagram of the Geldart classification of particles (Geldart, 1973)
4.2.1.3 Bubbling fluidized beds
Most operated fluidized beds are in the bubbling bed regime (Figure 4.3,C). This type of fluidization
is also called aggregative fluidization and under these conditions, the bed appears to be divided into
Technische Universiteit Eindhoven University of Technology
42 Evaluation of reactor concepts for ATR
two phases, the bubble phase and the emulsion phase. The behavior of these two phases are strongly
dependent on the bubbles and their properties. Several important properties include:
- The minimal fluidization velocity (which is already discussed in chapter 3.1.3)
- Bubble size
- Bubble wake
- Bubble rise velocity
- Flow pattern
- Bed expansion
4.2.1.4 Advantages and disadvantages of fluidized beds
Fluidized beds offer several advantages as compared to other reactor types. These can be
summarized as follows (Deshmukh et al., 2007; Kunii and Levenspiel, 1991):
- negligible pressure drop, small particle sizes can be used which result in no internal mass
and heat transfer limitations;
- excellent mixing properties due to the fluid-like behavior of the solid material;
- the possibility to operate isothermal;
- a great degree of freedom and flexibility in membrane and heat transfer surface area
placement;
- ability to operate in continuous state;
- Additionally, the membranes (if chosen well) improve fluidization behavior due to:
o Compartmentalization, reduced gas back mixing
o Reduced average bubble size because of enhanced bubble breakage, improving the
bubble to emulsion mass transfer.
The fluidized bed also has some disadvantages:
- a larger reactor vessel will be needed compared to a packed bed due to the expansion of the
bed at high gas velocities
- large feed flows of gas are required for fluidization
- particle entrainment could occur with very large flows
- The fluidization of particles could cause erosion
- pressure loss scenario in which the fluidization is suddenly lost. This could cause runaway
reactions or even dangerous situations (with highly exothermic reactions)
4.2.2 Preparation and procedure
The setup used is mainly designed for steam methane reforming and water gas shift, but
autothermal reforming of methane and oxidative coupling of methane for instance could also be
performed. First, an overview of the setup will be given after which the procedure followed for the
tests are elaborated on.
Technische Universiteit Eindhoven University of Technology
43 Evaluation of reactor concepts for ATR
4.2.2.1 FBMR setup
Figure 4.5 Schematic overview of the setup in which the FBMR is tested
The setup can be split in a few sections. First there is the feed section which includes the possibility
to feed several gases at a controlled flow rate due to the mass flow controllers. With the gases
available, reactions like steam methane reforming, water gas shift and oxidative coupling of methane
can be performed. There is also a possibility to feed steam or ethanol through the controlled
evaporation module (CEM). The lines and parts of the setup which could possibly come in contact
with steam are provided with a tracing, keeping the temperature of the lines high enough to prevent
steam from condensing.
The next section is the reactor. The reactor itself is 10 cm in diameter and 62 cm in height. The
distributor plate has a pore size of 40 µm which distributes the gas to the bed evenly and prevents
particles from falling through the distributor plate. When performing an experiment with the FBMR,
five membranes are mounted in the reactor. The membranes have an effective height of about 10
cm and a diameter of about 1 cm. In the reactor it would then look like presented in Figure 4.6. The
membranes are connected at the top to a space where all five lines connect to, resulting in one
permeate stream. It is also possible to bypass the reactor.
Figure 4.6 schematic of the inside of the fluidized bed membrane reactor with five tubular membranes installed via the Swagelok
graphite ferrules sealing method
The permeate stream can be collected using vacuum and is sent to the vent or the analyzer. For
prevention, water trap has been installed also at the permeate stream to prevent steam from entering
the analyzer. The analyzer, a SICK GMS810, is able to detect hydrogen (0 – 100 vol%), CO2 (0 – 100
ppm) and CO (0 – 200 ppm).
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44 Evaluation of reactor concepts for ATR
The retentate stream goes through two coolers before either being sent to the vent or the analyzer.
A filter installed with a pore size of 15 µm which capture any possible entrained particles. The
analyzer for the retentate stream is a SICK GMS815P which is able to detect CO, CO2, CH4 and H2
(0 – 50 vol%).
Safety features are installed, ensuring safe working conditions. When the reactor pressure exceeds
6 bar, the safety valve will open releasing all excess pressure to the vent. Next to this physical form
of safety feature, the software is also included with safety features.
4.2.2.2 Procedure
The same procedure has been followed for all tests using the FBMR which includes that the
membrane selectivity and performance is first checked before performing any reaction. This means
the membranes are first installed in the reactor without any particles and tests are performed with
nitrogen to evaluate the leakage, hydrogen to evaluate the permeability and possibly mixed with CO
to evaluate the selectivity of the membrane compared to CO.
The leakage is evaluated at room temperature for nitrogen, and at high temperature (usually around
500 °C) for all components. The membranes were heated up using a ramp of 5 °C/min under a flow
of nitrogen.
If the preliminary results are within the preset ranges (selectivity > 10,000), then the catalyst was
added. In this case 0.5 kg of 2 wt% Ru/CeZrO2 mixed with 3.5 kg of zirconia based filler. 4 kg of
material is used to ensure full submersion of the membranes. This hopefully results in higher
hydrogen yields and higher methane conversions.
The reactor is heated up under fluidization conditions determined by using the correlation of Wen
and Yu as explained in Chapter 3.1.3. It was made sure that the actual gas velocity was always above
the 1.5 umf during heat up to prevent the bed acting as a fixed bed. The bed was heated up with a
ramp of 2 °C/min. It should be noted that at lower temperatures (T < 150 °C) the bed was not
fluidized due to limitations of the setup in terms of gas that could be supplied.
At the desired temperature, the catalyst was first activated using a diluted stream of hydrogen. A 20
vol% hydrogen stream was fed for 2 hours to activate the catalyst. This is also beneficial for the
membranes as they have to be activated as well.
After activation the desired reaction was performed. This often included first steam methane
reforming as the membrane sealings could be prone to oxidation. So first SMR was performed
instead of ATR. Different configurations were tested, the steam-to-carbon ratio and also the amount
of dilution.
Technische Universiteit Eindhoven University of Technology
45 Evaluation of reactor concepts for ATR
4.2.3 Results
The results of the tests in the FBMR are presented in two sections, first the test with commercially
available membranes will be discussed and later the test with the Tecnalia membranes will be
discussed.
To analyze the data, some parameters are defined to quantify the performance:
Methane conversion 4 4
4
, ,
,
CH in CH out
CH in
4.1
H2/CO ratio 2 ,ret
,
H
CO ret
4.2
CO selectivity
2
,
, ,
CO ret
CO ret CO ret
4.3
H2 selectivity 2
4 2
,
, ,1( ) 42
H total
CH reacted O in
4.4
Hydrogen recovery factor 2
4 2
,
1( ),in2
4H perm
CH O
HRF
4.5
Separation factor 2
2
,permeated
,total
H
H
SF
4.6
4.2.3.1 FBMR with commercial membranes (REB)
For the test with the REB membranes, autothermal reforming of methane was evaluated. The REB
membranes were used to evaluate the performance of the catalyst in a FBMR. The membranes are
known to be stable with a high selectivity but a lower permeability compared to the new Tecnalia
membranes. The concept of autothermal reforming in a FBMR could therefore be evaluated quickly.
The test was performed for three days to check the stability of the system. Figure 4.7 gives an
overview of the second day where the reactor pressure was increased to see the effect of pressure.
Figure 4.7 Autothermal reforming (ATR) in a FBMR with REB commercial hydrogen membranes at 600 °C at 1.3 and 1.47 bar.
Feed: SCR of 1.91, OCR of 0.43 and an NCR of 5.24 with a total flow rate of 11.25 Nl/min.
0 50 1000
5
10
15
20
25
30
CO CO2 CH4 H2 CH4 conversion (%)
Ret
enta
te c
ompo
sitio
n (v
ol%
)
Time (min)
0
50
100
Reaction and separation600 °C 1.47 bar
Reaction 600 °C 1.47 bar
Reaction and separation600 °C 1.3 bar
Reaction 600 °C 1.3 bar
CH
4 co
nver
sion
(%)
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46 Evaluation of reactor concepts for ATR
Table 4.2 Summary of the FBMR test performed with the REB membranes at 600 °C at 1.3 and 1.47 bar
1.3 bar 1.47 bar
CH4 eq. conversion (%) 88.8 87.2
u/umf 1.5 1.3
Membrane integration No Yes No Yes
CH4 conversion (%) 90.4 92.9 89.4 92.9
H2/CO ratio 8.3 7.3 8.6 7.5
CO selectivity 0.33 0.29 0.32 0.28
H2 selectivity 0.88 0.89 0.89 0.91
HRF 0.20 0.21
SF 0.24 0.26
From Table 4.2 it can be concluded that the autothermal reforming of methane can be performed in
a FBMR. The hydrogen recovery factor (HRF), which is determined as the hydrogen recovered by
the membrane compared to the theoretical maximum production of hydrogen, was 20% and 21%
respectively.
The separation factor indicates how much hydrogen that has been produced permeated through the
membrane. Roughly 25% of the produced hydrogen is permeated and can be classified as ultrapure
hydrogen. This is still a low percentage and should be improved, the easiest way is to increase the
pressure in the reactor. Due to limitations to the setup, this was not possible for this experiment.
With increasing pressure, it is expected that the permeation through the membrane is higher due to
a higher transmembrane pressure difference. The increase in permeate flow was expected to be
6.5%. The observed value of 4.5% is probably lower due to limitations of the system in terms of
concentration polarization.
Figure 4.8 Autothermal reforming (ATR) in a FBMR with REB commercial hydrogen membranes at 600 °C at 1.48 and 1.58
bar. Feed: SCR of 1.49, OCR of 0.43 and an NCR of 4.44 with a total flow rate of 12.4 Nl/min and SCR of 1.91, OCR of 0.43 and
an NCR of 5.3 with a total flow rate of 11.25 Nl/min.
Table 4.3 Summary of the FBMR test performed with the REB membranes at 600 °C at 1.48 and 1.58 bar with different feed
flowrates.
1.48 bar / high flowrate 1.58 bar / low flowrate
CH4 eq. conversion (%) 87.2 86.4
u/umf 1.5 1.2
Membrane integration No Yes No Yes
CH4 conversion (%) 85.9 87.4 88.7 92.3
H2/CO ratio 7.17 6.13 8.41 7.26
CO selectivity 0.38 0.35 0.33 0.29
H2 selectivity 0.89 0.90 0.89 0.89
HRF 0.18 0.20
SF 0.23 0.26
0 50 100 1500
5
10
15
20
25
30
Reaction and separation 600 °C 1.58 bar
Reaction 600 °C 1.58 bar
Reaction and separation 600 °C 1.48 barhigher flowrate
Reaction 600 °C 1.48 barhigher flowrate
Ret
enta
te c
ompo
sitio
n (v
ol%
)
Time (min)
CO CO2 CH4 H2 CH4 conversion (%)
0
50
100
CH
4 co
nver
sion
(%)
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47 Evaluation of reactor concepts for ATR
The effect of the flowrate has also been investigated. The flowrate inherently influences the behavior
of the bed. From the results in Table 4.3 it can be concluded that a low flowrate is more beneficial
for the system in terms of methane conversion and separation factor. Especially in the methane
conversion it can be concluded that a longer residence time is beneficial for the system to extract
hydrogen and to shift the equilibrium of the system.
Overall it can be concluded that the concept of ATR in a fluidized bed is viable, although the effect
of shifting the equilibrium by the use of membranes is smaller than expected. If the case of 1.3 bar
is considered (Figure 4.7), in the equilibrium state the methane conversion is about 90% and and
the HRF is 0.20. The overall hydrogen production due to implementation of the membranes has
increased with only 2.1%.
Shifting the equilibrium towards hydrogen due to integration of the membranes has some room for
improvement. Hydrogen extracted compared to the hydrogen produced (without integrating the
membranes) is 20.5%, which is ultrapure. It was unfortunately not possible to monitor the purity of
the permeate stream. The REB membranes were also used by Patil and he reported a selectivity of >
10,000. It is assumed the membranes still hold this selectivity.
4.2.3.2 FBMR with Tecnalia tubular membranes
In the first attempt to perform experiments in the FBMR with the Tecnalia membranes, the
preliminary testing showed a membrane selectivity of >10,000 so the catalyst was added and the
reactor was heated up. Heating up was not performed in inert atmosphere, there was 5 vol% oxygen
present with the intention to burn any impurities. This could have caused the observed leakage after
heating up, dropping the selectivity far below 10,000. It was identified the leakage occurs at the
sealings. Later, the sealing has been investigated for different gases and temperatures to evaluate its
stability (Chapter 4.3).
Because of the low selectivity of the membranes, the reactor had to be cooled down, opened and the
membranes had to be examined to improve the sealing to reach the required selectivity again. In the
third try, reaction could be performed and the results are discussed below.
SMR reaction was performed first to exclude effects of oxygen on the system. From the previous test
on the membrane it was known, that until 500 °C, the membrane sealings are stable. After 500 °C
the sealings are unstable and the nitrogen leakage starts to increase. So it was decided to start testing
at 500 °C and stepwise increase the temperature to 600 °C. Graphs of the experiments with the
Tecnalia membranes can be found in the appendix.
Table 4.4 Summary of SMR in the FBMR with Tecnalia tubular Pd-Ag/ZrO2 membranes at 1.3 bar at several temperatures.
SMR in FBMR (Tecnalia membranes) 500 °C 550 °C (day 1) 550 °C (day 2) 600 °C
CH4 eq. conversion (%) 55.7 73.0 73.0 88.1
u/umf 1.3 1.3 1.3 1.5
CH4 conversion (%) 55.5 73.1 76.4 89.3
H2/CO ratio 22.6 16.1 15.8 11.0
CO selectivity 0.12 0.16 0.18 0.25
H2 selectivity 0.96 0.96 0.96 0.94
HRF 0.17 0.22 0.20 0.23
SF 0.31 0.31 0.28 0.28
H2 impurity (ppm CO) 50 70 120 200
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48 Evaluation of reactor concepts for ATR
Some trends which we can see from the data in Table 4.4 is that with increasing temperature, the
methane conversion is higher. This makes sense as the steam methane reforming reaction is
endothermic. The water gas shift, which follows, is exothermic, but the CO needed for WGS is
produced in the SMR reaction. At lower temperatures, less CO is observed. This is expected as the
WGS equilibrium towards hydrogen is more favorable at lower temperatures.
From the membrane study, the Tecnalia membranes are expected to perform better in terms of
hydrogen permeation than the REB membranes. The separation factor (SF) is an important
parameter to compare the different membranes. In the test with the REB membranes a higher
transmembrane hydrogen pressure was used but the SF of the Tecnalia membranes are significantly
better. If the hydrogen pressure in the Tecnalia test would be the same as in the REB test, the SF
would even be higher, so it can be concluded Tecnalia membranes are better in terms of hydrogen
permeability in the FBMR.
The autothermal reforming of methane has also been investigated. At this point it was already
impossible to analyze the permeate stream because the CO concentration was out of range for the
analyzer (and the selectivity of the membranes was thus <<10.000).
Table 4.5 Comparison of ATR/SMR in the FBMR with REB and Tecnalia membranes at 600 °C.
REB
membranes
ATR Tecnalia
membranes
SMR Tecnalia
membranes
CH4 eq. conversion (%) 88.8 93.2 88.1
u/umf 1.3 1.3 1.5
CH4 conversion (%) 92.9 96.7 89.3
H2/CO ratio 7.3 10.4 11.0
CO selectivity 0.29 0.22 0.25
H2 selectivity 0.89 0.97 0.94
HRF 0.20 0.35 0.23
SF 0.24 0.31 0.28
It has to be noted that the feed composition was different. The SCR is 1.9 with the REB membranes
versus 3.0 with the Tecnalia membranes. It is evident the Tecnalia membranes are performing better
based on the separation factor. It can also be observed that the H2 and CO selectivities are higher,
possibly due to the higher residence time. Due to the high content of CO in the permeate, the
permeate stream was not analyzed.
These tests have shown that SMR and ATR can be performed in the FBMR with Tecnalia tubular
membranes. The selectivity of the membrane, or rather the leakage of the sealings, is an important
topic which will be discussed in the Chapter 4.3. This has major impact on the purity of the hydrogen
permeated and the aim is ultrapure. The permeation of the Tecnalia membranes compared to the
REB membranes in similar systems is significantly better but still have room for improvement.
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49 Evaluation of reactor concepts for ATR
4.3 Membrane stability in fluidized beds
The tests with the FBMR showed that the leakage of the membranes was getting higher during the
test. Separate tests on membranes have been conducted to find the reason for the observed leakage.
In addition, several protections of the sealings were applied to evaluate the contribution of particle
interaction with the sealing.
4.3.1 Preparation of sealing protections
Several sealed membranes were prepared with different kinds of protections to prevent any particles
from reaching the sealing. The first protections that were applied were applying a paste around the
sealing. Two different pastes were used which are able to withstand high temperatures.
Figure 4.9 Two membranes with paste at the sealings to prevent particles from interacting with the graphite gasket, black insulation
paste (l) and Ceramabond by Aremco (r)
The graphite ferrules are installed in the Swagelok connections with the flat part to the open part of
the connection. The graphite gasket also has a more pointy part and by mirroring the graphite gasket
the exposed area of the graphite is significantly reduced. Next to reversing the sealings, a membrane
was also prepared with an extra insulation ring before the graphite gaskets.
The membrane was then placed in a fluidized bed reactor and under fluidization conditions, the
leakage over time and temperature was monitored.
4.3.2 Results
The same procedure had been followed as for the FBMR to decide if the test would be conducted.
This means a selectivity of over 10,000 was required to start measuring. The two sealings protected
by high temperature paste proved to be no solution. After heating up the membrane, the leakage was
significantly higher which made testing useless. Either the paste started interacting with the
membrane surface or the mechanical properties of the paste (like thermal expansion) could not be
handled by the membrane. It was decided not to investigate why it started leaking as we were more
interested in trying to find a solution.
The tests with the reversed ferrules were successful. As a reference, a base case was also performed
where the membrane had the sealings in the conventional way and no protection in any form was
applied.
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50 Evaluation of reactor concepts for ATR
Figure 4.10 Nitrogen leakage and temperature of the reactor over time for conventional ferrule placement (top left), reversed
ferrule placement (top right) and reversed ferrules placement + extra protection ring (bottom)
The trend that can be observed is that with increasing temperature, nitrogen leakage decreases. This
is expected, since the graphite ferrules expand with increasing temperature, squeezing in the
Swagelok connection. This sealing in terms of nitrogen leakage is improving until about 500 °C.
This is also the temperature the supplier gives as the maximum operational temperature (in reactive
conditions). After 500 °C a steep increase in leakage is observed with the reversed ferrules, with or
without protection ring. The membrane with the conventional placed ferrule doesn’t show the same
increase of leakage. This membrane had already been used for previous testing and had already been
subjected to high temperatures. In Chapter 2.3.2.2, a long term test was also performed at high
temperature but without fluidizing particles. If the leakage is normalized from the value at 500 °C,
the following graph is obtained.
Figure 4.11 Normalized nitrogen leakage after 500 °C for the conventional placed ferrules without fluidization conditions, reversed
and reversed + protected ferrule placement under fluidization conditions
0 100 200
0
100
200
300
400
500
600
Tem
pera
ture
(°C
)
Time (h)
0.0
1.0x10-12
2.0x10-12
3.0x10-12
Nitr
ogen
leak
age
(mol
s-1 P
a-1)
0 100 2000
100
200
300
400
500
600
Tem
pera
ture
(°C
)
Time (h)
0.0
5.0x10-12
1.0x10-11
Nitr
ogen
leak
age
(mol
s-1 P
a-1)
0 20 40 60 80 100 120 140 160 1800
100
200
300
400
500
600
Tem
pera
ture
(°C
)
Time (h)
0.0
2.0x10-12
4.0x10-12
6.0x10-12
Nitr
ogen
leak
age
(mol
s-1 P
a-1)
0 20 40 60 80 100 120
1
2
3
4
5
6
Conventional Reversed Reversed + protected
Nitr
ogen
leak
age
norm
aliz
ed (-
)
Time (h)
Technische Universiteit Eindhoven University of Technology
51 Evaluation of reactor concepts for ATR
Independent on fluidization conditions or not, it is clear the leakage increases at high temperature
over time. Fluidization conditions, however, seem to accelerate this process of increase in leakage.
It can be conclude that at high temperature (T > 500 °C) the ferrules cannot hold and the leakage
starts to increase. Next to the effect of temperature, the effect of several gases has been determined.
Figure 4.12 Nitrogen leakage after exposing the membrane to several gases at different temperatures and durations
From the data presented in Figure 4.12, it seems that there is no significant effect from CO, CO2 and
steam. The increase after 15 hours of steam looks significant. To confirm it’s indeed the effect of
steam and not of temperature, the membrane was exposed to nitrogen for 15 hours after which a
similar relative increase was found. Therefore it’s more probable the cause of the increase in leakage
is the temperature.
After the tests, the membranes were visually inspected to check if they suffer from defects. The
membranes were also tested in water/ethanol to identify in which part of the membrane the leakage
is present. From these tests it can be concluded that the majority of the leakage is caused at the
sealing of the membrane. With some membranes there were also pinholes observed, but compared
to the leakage at the sealings, this is negligible. An impression of the test can be found in Figure
4.13.
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52 Evaluation of reactor concepts for ATR
Figure 4.13 image of a membrane used in the FBMR tested in water with air at 1 barg on the inside of the membrane
Some membranes which have been exposed to the Ru/CeZrO2 catalyst have also been analyzed
using SEM-EDX to determine if any surface reactions between catalyst and membranes have been
present. The full analysis was done by Tecnalia. For the analysis the membranes were first examined
under an optical microscope. Based on the images there, the membranes could be divided in three
categories.
Table 4.6 Classification given to membranes based on optical microscopy
Membrane # Category Membrane surface appearance 1 I High surface roughness and some sheets 2,5,7 II High amount of particles 3,4,6 III Low amount of particles
Of every category one membrane was investigated with the SEM-EDX. The membrane had been
divided into four main areas of interest which can be found in Figure 4.14. The four areas of interest
are the central area of the membrane (a), the area near the membrane-graphite zone where the
stagnant particles were (b), the interphase between the membrane and the graphite (c) and the
graphite zone (d).
Figure 4.14 membrane with the depicted areas analyzed by SEM-EDX
SEM images were taken to have a closer look on the membrane. The membranes of category II are
particularly interesting. From optical microscopy a lot of particles could be seen, together with EDX
it’s possible to retrieve the composition of these particles.
d c b a
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53 Evaluation of reactor concepts for ATR
Figure 4.15 SEM pictures of a membrane classified as having a high amount of particles on the surface at the central area (a) (top
left and right) the interphase between membrane and graphite (c) (bottom left) and the graphite zone (d) (bottom right)
In these SEM images the particles are really distinguishable. It can be seen that in Figure 4.15a the
size of the particle seems to be around 15 µm and it’s quite interesting to know what the composition
is.
Figure 4.16 mapping surface of the central zone of a membrane classified as having a high amount of particles by EDX
The EDX of the central zone (a) shows the expected Pd and Ag but also a high amount of Ru is found.
This indicates the particles which are found on the surface are ruthenium particles. There was also
a small amount of ceria found and no zirconia, which indicates that catalyst particles broke up and
the ruthenium and ceria showed interaction with the Pd-Ag membrane surface. The zirconia, and
also the zirconia based filler, are not detected in the EDX analysis of the membranes so it’s safe to
conclude that the filler shows no interaction with the Pd-Ag membrane surface.
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54 Evaluation of reactor concepts for ATR
Table 4.7 Summary of the EDX results of the membrane with high amount of particles on the surface (composition in wt%)
Spectrum O Zr Ru Pd Ag Ce Total
Mapping of “a” (central area) 200x 7.36 74.95 13.36 1.86 2.48 100.00
Dark particle in “a” (central area) 15000x 73.35 25.41 1.24 100.00
Deep background near the dark particle in “a” (central area) 15000x 3.87 31.14 55.83 7.60 1.56 100.00
Background in “d” (graphite zone) 5 mm far from the interphase 15000x 7.57 76.70 10.45 1.10 4.17 100.00
The ruthenium particles are mainly picked up by EDX. It can also be seen that when the focus is on
a part with no visible particle, the amount of ruthenium is significantly less. It’s also quite interesting
to see that in zone d, where the membrane was touching the graphite sealing, the same amount of
ruthenium is found as in the central zone. This is most probably because the membrane had
probably been used before and had been cut and resealed.
Figure 4.17 SEM image (l) and EDX analysis (r) on the membrane classified as high surface roughness and some sheets
It is also quite interesting to compare the results with the membrane where no particles were found
on the surface in the central zone (a) to see if there is a difference, these results can be found in
Figure 4.17. From the SEM image it can be seen that no particles are present and that the surface is
indeed rough. The EDX analysis is almost identical as the EDX analysis with the membrane with a
high amount of visible particles. Unfortunately there is no higher magnification picture but there
are small particles visible which, in contrast to the other membrane, look like they crashed on the
surface, hence probably also the reason no craters are visible but shatter cones, and were dragged
over the surface.
It’s clear that ruthenium shows interaction with the Pd-Ag surface of the membrane . The filler has
no apparent interaction with the membrane surface. These results, however, do not give an
explanation of the increased leakage. The EDX analysis of the central zone (a) and the interphase of
the membrane with the graphite (c) look similar so that cannot explain the leakage which is observed
at the sealings.
If we summarize the results on the nitrogen leakage, temperature itself is a cause for the increase in
leakage. Fluidization conditions seem to make the increase in nitrogen leakage higher but this could
also be the reversing of the ferrule. Other single gases do not show any additional significant increase
of the leakage. Visual tests show that the leakage is at the sealings and not at the membrane surface.
EDX analysis does not show a significant different composition close to the sealings.
Technische Universiteit Eindhoven University of Technology
55 Evaluation of reactor concepts for ATR
4.4 Conclusion
It can be concluded that autothermal reforming of methane in a microreactor as performed in this
work is not viable. In the case of the wall-coated microreactors, no hydrogen yield was observed. In
the case of the packed bed (membrane) microreactor, hydrogen was observed, but in quantities much
lower than expected. Because of the low hydrogen yield and therefore a low hydrogen partial
pressure, the hydrogen permeation in the PBMMR was non-existent.
The fluidized bed membrane reactor showed that autothermal reforming of methane can be
performed in a FBMR. High methane conversions can be achieved which are close to the
equilibrium of the system. The REB membranes could subtract ultrapure hydrogen with a SF of 0.2
so the system is able to produce ultrapure hydrogen, although in low quantities, without the need of
additional separation units. The second purpose of introducing membranes is to improve the
methane conversion and produce additional hydrogen due to shifting the equilibrium. This effect
however was almost non-existent with a mere 2% additional hydrogen produced.
The FBMR system with the Tecnalia membranes showed to be unstable. The sealings of the
membranes start to leak at high temperatures. In the performed tests, the system was able to reach
equilibrium conversion and the Tecnalia membranes show a better separation factor under worse
conditions which indicates the Tecnalia membranes are significantly better than the REB
membranes.
The catalyst deactivation which has been reported in Chapter 3.2.3.2 is not detectable in the fluidized
bed. This is most probably because the amount of catalyst used in the FBMR is that high that
deactivation will probably never be detected due to the excess of catalyst.
It can be concluded that both reactor concepts experimentally are quite devious. Therefore a
theoretical comparison will be made in Chapter 5. In this case, the reactor concepts will be compared
at ideal conditions.
A close look has been given to the leaking of the Tecnalia membranes. Leakage occurs at the sealings
at temperatures over 500 °C and fluidization conditions accelerate the slope of the leakage.
Ruthenium and in less extent ceria show interaction with the Pd-Ag surface of the membrane. The
zirconia based filler does not show any interaction with the Pd-Ag surface of the membrane. From
this data, the most probably cause for leakage is temperature alone.
Technische Universiteit Eindhoven University of Technology
56 Theoretical comparison reactor types for ATR of methane
5. Theoretical comparison reactor types for ATR
of methane
The catalyst and membrane separately look promising for the ATR of methane, integration of the
catalyst and membranes was not completely successful. Integration of the catalyst with the
membranes to produce ultrapure hydrogen is essential and the choice of reactor type is equally
important. A theoretical comparison is made between the Packed Bed Membrane Microreactor
(PBMMR) and the Fluidized Bed Membrane Reactor (FBMR) in which advantages and
disadvantages of the reactor concepts regarding the ATR of methane will be evaluated. For the
comparison it is assumed that catalyst and membrane are the same so solely the reactor types will
be compared.
For the comparison, the two reactor types have been modelled using models available within the
SMR group to compare the performances of the reactor types. Economic aspects will also be taken
into account to compare the preliminary costs of a small scale production plant with these reactor
types.
The PBMMR has been modelled using an available code for packed bed reactors. The code has been
adjusted to be modeled for a microchannel. The model is using a different geometry though, a
cylindrical geometry, whereas the microchannels used experimentally were squared. Additionally,
the membranes are modelled as an internal tube.
The FBMR has been modelled as a vessel with a certain amount of effective membrane area. The
effective membrane area is determined to be similar as used in the experimental setup.
Both the PBMMR and the FBMR have been modelled using the same kinetics, catalyst particles,
membrane permeability and thickness so the comparison is based on reactor type. The inlet
composition was determined as SCR = 1.242, OCR = 0.389 and NCR = 1. This is based on the work
of Patil where these values for the SCR and OCR resulted in autothermal operation (Patil, 2005).
Technische Universiteit Eindhoven University of Technology
57 Theoretical comparison reactor types for ATR of methane
5.1 Reactor concepts in literature
Both reactor types, the PBMMR and the FBMR, for hydrogen production have recently caught
attention. The novel reactor configurations are being investigated due to their many advantages over
the more conventional packed bed membrane reactor. A summary of interesting findings in
literature has been summarized and given by (Gallucci et al., 2013).
The packed bed membrane microreactor, is particularly performing better than the packed bed
membrane reactor due to:
- Mass and heat transfer is improved because of the scale length in microchannels
- Concentration polarization is negligible
(A. L. Mejdell et al., 2009b) showed that the concentration polarization can be neglected in a
membrane microreactor whereas concentration polarization is the limiting step for hydrogen
permeation in a tubular configuration.
The fluidized bed concept has already been discussed in Chapter 4.2.
Another finding in literature is the effectiveness factor for hydrogen permeation proposed by (CHEN
et al., 2007) and (Mahecha-Botero et al., 2008) which is an adaptation to the Sieverts’ law for
hydrogen permeation. The effectiveness factor proposed is 0.9. This has not been taken into account
in the model.
When comparing reactor concepts, important parameters to look at are:
- Extent of mass transfer limitations
- Heat supply to or heat removal from the reactor
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58 Theoretical comparison reactor types for ATR of methane
Modelling
Packed bed membrane microreactor
The PBMMR has been modeled using a 1D packed bed model which was already available within the SMR group. It is pseudo homogeneous and using the implicit particle model. It is modeled as a tube in a tube where the inner tube is the permeate and the outer tube is the retentate as depicted in Figure 5.1. Kinetics are taken from (Numaguchi and Kikuchi, 1988) for the SMR and WGS reactions. For the oxidation of methane, the kinetics of (Trimm and Lam, 1980) are used. The permeation constants determined in chapter 2.3.2 for the Pd-Ag/ZrO2 tubular membrane of Tecnalia are used in the model for the permeation of hydrogen.
Figure 5.1 Schematic representation of the PBMMR model
Some key parameters that have been used for the simulation are listed in Table 5.1.
Table 5.1 parameters and constants used for the modelling of ATR in a PBMMR
PBMMR Reactor diameter (m) 0.005 Membrane diameter (m) 0.001 # of membranes 1 Length of the bed (m) 0.03 – 0.10 Pressure reactor (bar) 5 – 11 Pressure retentate (bar) 0.03 Catalyst in bed (%) 100 Membrane area per reactor volume (m2
membrane/m3reactor)
166.6
Feed flow rate (mol/m2/s) 12.3 Temperature of feed (°C) 600
The feed composition was taken from (Patil, 2005) where it was concluded that an SCR of 1.242 and an OCR of 0.389 would result in autothermal conditions. Reactor dimensions were chosen to be viable in industry. First the length of the reactor has been evaluated for this system. Some simulations have been run without membrane integration to validate the model is converging to the equilibrium conversion and this was compared to either isothermal or adiabatic systems with membrane integration. In Figure 5.2 it can be seen the base case is converging to the equilibrium conversion as expected. With membrane integration but running at isothermal conditions, it can be seen that the methane conversion can reach almost 100%. It can also be seen that the integration of membranes has a minor effect on the methane conversion in the first part of the reactor.
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59 Theoretical comparison reactor types for ATR of methane
Figure 5.2 comparison simulation results between the base case without membrane integration and with the isothermal case
To see the effect of the steam-to-carbon ratio, some simulations have been run without membrane
integration at isothermal conditions. Inherently the equilibrium conversion changes when adjusting
the concentrations as the feed flow rate was kept the same. From Figure 5.3 (l) it can be seen that the
reaction rate is almost similar for all three cases. In terms of reaching the equilibrium conversion, a
lower steam to carbon ratio is beneficial. This is due to the inlet feed which contains more methane
which is more important for the reaction rate of steam methane reforming (Numaguchi and Kikuchi,
1988).
When integrating the membranes, a lower steam to carbon ratio seems to be favored. In Figure 5.3
(r) it can be seen that not only the overall methane conversion is higher at lower SCR but also the
reaction rate is higher at lower SCR due to the higher methane concentration present. From the
results it can be concluded that there is no reason to use higher SCR’s, not in terms of methane
conversion.
Figure 5.3 Comparison of different SCR ratios without membrane integration (l) and with membrane integration (r)
The next step was to run the simulations adiabatically. The simulation was run adiabatically to
evaluate the temperature effect. In Figure 5.4 it can be seen that the reaction rate is significantly
higher in the first part of the reactor due to the high temperature. The methane conversion also
drops halfway due to the temperature decrease and the changed thermodynamic equilibrium. It is
important the temperature is controllable for the stability of the membranes.
0.00 0.02 0.04 0.06 0.08 0.100
20
40
60
80
100
Met
hane
con
vers
ion
(%)
Reactor length (m)
No membrane Isothermal
Eq. conversion
0.00 0.02 0.04 0.06 0.08 0.100%
20%
40%
60%
80%
100%
SCR 1.242 SCR 1.9 SCR 2.5
Met
hane
con
vers
ion
z (m)
Actual methane conversion Normalized methane conversion (X/Xeq)
0.00 0.02 0.04 0.06 0.08 0.100%
20%
40%
60%
80%
100%
Met
hane
con
vers
ion
z (m)
SCR 1.242 SCR 1.9 SCR 2.5
Actual conversion Equilibrium conversion
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60 Theoretical comparison reactor types for ATR of methane
Figure 5.4 Comparison methane conversion in different systems (l) and the temperature profile of the adiabatic simulation (r)
In Figure 5.5 the results of the simulations where cooling is applied are shown. With different cooling
temperatures and heat transfer coefficients it proves difficult to cool the PBMMR in such a way, a
uniform temperature is achieved. The oxidation of methane in the first part of the reactor is a big
problem in terms of temperature distribution and cooling doesn’t show the desired effect. In
practice, either local cooling needs to be applied or the heat transfer should be increased for instance
by introducing cooling coils in the reactor, which in its turn is nearly impossible at the used reactor
dimensions.
Figure 5.5 Comparison methane conversion with cooling (l) and the accompanying temperature profiles (r)
For the comparison of the simulations with the FBMR, the isothermal case will be considered to be
able to compare the intrinsic reactor types.
5.2.2 Fluidized bed membrane reactor
The FBMR has been modeled using a one dimensional two-phase model which divides the fluidized
bed in stirred tank reactors in series. A schematic representation is added as Figure 5.6. The kinetics
and membrane constants are the same as used in the PBMMR model.
Figure 5.6 Schematic representation of the FBMR model (Patil, 2005)
0.00 0.02 0.04 0.06 0.08 0.100
20
40
60
80
100
Met
hane
con
vers
ion
(%)
Reactor length (m)
No membrane integration Isothermal Adiabatic Equilibrium conversion
SCR = 1.242OCR = 0.379T_in = 873 K
0.00 0.02 0.04 0.06 0.08 0.10800
850
900
950
1000
1050
1100
1150
1200
Tem
pera
ture
(K)
z (m)
Tw_1 Tg_1 Tw_2 Tg_2
0.00 0.020%
20%
40%
60%
80%
100%
met
hane
con
vers
ion
z (m)
Adiabatic 300 10 W/m2K 500 10 W/m2K 500 100 W/m2K
0.00 0.02600
650
700
750
800
850
900
950
1000
1050
1100
1150
1200
Tem
pera
ture
(K)
z (m)
Adiabatic 300 10 W/m2K 500 10 W/m2K 500 100 W/m2K
Tw_1 Tg_1
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61 Theoretical comparison reactor types for ATR of methane
Some key parameters that have been used for the simulation are listed in Table 5.2. The attempt has
been made to keep some parameters the same as in the PBMMR to be able to compare the results
later. This means for instance a high membrane area per reactor volume is assumed.
Table 5.2 parameters and constants used for the modelling of ATR in a FBMR
FBMR
Reactor diameter (m) 0.42
Membrane diameter (m) 0.01
# of membranes 200 - 500
Length of the bed (m) 0.20
Pressure reactor (bar) 8
Pressure retentate (bar) 0.03
Catalyst in bed (%) 20 - 50
Membrane area per reactor volume
(m2membrane/m3
reactor)
91 - 272
Feed flow rate (mol/m2/s) 10.9
u/umf 5 - 2.5
Temperature of feed (°C) 600
First, a preliminary test was done to see if the code converges to the equilibrium conversion. In
Figure 5.7 it can be seen this is indeed the case. Membrane integration has a similar effect as for the
PBMMR and in the FBMR it is also possible to reach high conversions of methane. The main
difference observed in the adiabatic simulation is that in the FBMR the conversion is the same as in
the isothermal case. This is because there is no temperature profile within the bed and the inlet
concentrations as proposed by Patil are used, so no temperature increase was expected and observed.
Figure 5.7 comparison methane conversion in the FBMR
From theory and experiments it is expected that catalyst particles are abundant in this system. Some
simulations have been done to confirm if the amount of catalyst is indeed abundant. From Figure
5.8 it seems the catalyst is abundant in the system as the methane conversion isn’t increasing or
decreasing too much. This is also confirmed by the numbers as shown in Table 5.3. The hydrogen
recovery factor can only be increased by 22% when the catalyst loading is multiplied by 2.5.
0.0 0.1 0.2 0.3 0.4 0.5 0.6 0.70%
20%
40%
60%
80%
100%
Met
hane
con
vers
ion
z (m)
No membrane integration Isothermal Adiabatic Equilibrium conversion
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62 Theoretical comparison reactor types for ATR of methane
Figure 5.8 Comparison methane conversion with different
catalyst loadings in the bed
Table 5.3 Effect of catalyst loading on several indicators for ATR
Catalyst loading in bed
(%)
20 30 40 50
Methane conversion (%) 83.4 94.6 98.3 99.4
HRF (%) 74.2 88.7 93.8 95.8
SF (%) 95.4 97.2 98.4 98.9
CO selectivity 0.20 0.16 0.13 0.11
Permeate flow (Nm3/hr) 86 103 109 111
The membrane area was also varied to see the effect on the methane conversion. From Figure 5.9 it
can be seen that the effect of the membrane only starts when the methane conversion is at 40%.
This is also expected as the hydrogen partial pressure must be high enough for permeation to
happen. This also means that in practice, the membranes can be made shorter than the length of the
fluidized bed. Compared to the catalyst loading, the specific membrane area has a smaller effect on
the methane conversion and also on the permeate flow which is interesting but understandable due
to the driving force, hydrogen partial pressure, being smaller.
Figure 5.9 Comparison methane conversion with different
specific membrane areas
Table 5.4 Effect of specific membrane area on several indicators for ATR
Specific membrane area
(m2/m3)
91 136 181 272
Methane conversion (%) 87.2 94.6 97.0 98.2
HRF (%) 78.8 88.7 92.5 94.8
SF (%) 93.0 97.2 98.6 99.3
CO selectivity 0.20 0.16 0.13 0.09
Permeate flow (Nm3/hr) 91 103 107 110
0.00 0.02 0.04 0.06 0.08 0.10 0.12 0.14 0.16 0.18 0.200%
20%
40%
60%
80%
100%
Met
hane
con
vers
ion
z (m)
Catalyst in bed 20% 30% 40% 50%
0.00 0.02 0.04 0.06 0.08 0.10 0.12 0.14 0.16 0.18 0.200%
20%
40%
60%
80%
100%
Met
hane
con
vers
ion
z (m)
Specific membrane area (m2/m3) 91 136 181 272
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63 Theoretical comparison reactor types for ATR of methane
5.2.3 Comparison reactor types
To compare the performance of the two reactor types for the ATR of methane, the results of the
isothermal simulations are used to exclude temperature effects on the methane conversion. The
parameters used are shown in Table 5.5.
Table 5.5 parameters and constants used for the modelling of ATR in a PBMMR and FBMR
PBMMR FBMR
Reactor diameter (m) 0.005 0.42
Membrane diameter (m) 0.001 0.01
# of membranes 1 375
Length of the bed (m) 0.10 0.10
Pressure reactor (bar) 8 8
Pressure retentate (bar) 0.03 0.03
Catalyst in bed (%) 100 100
Membrane area per reactor volume
(m2membrane/m3
reactor)
166.6 166
Feed flow rate (mol/m2/s) 12.3 10.9
Temperature of feed (°C) 600 600
The methane conversion and hydrogen permeation of the two reactor models over the length of the
bed are shown in Figure 5.10. There is quite a big difference in the speed of reaction. The most
probable reason is the mass transfer limitations which are more present in the FBMR. The hydrogen
permeation in the FBMR is starting slower than in the PBMMR which is caused by the lower partial
pressure of hydrogen due to the lower conversion in the beginning of the bed.
Figure 5.10 Comparison of the FBMR and PBMMR model in terms of methane conversion and hydrogen permeation
The results are also presented in Table 5.6. Some clear differences can be observed. Although the
FBMR has a longer residence time, the PBMMR outperforms the FBMR in every aspect. The
apparent reaction rate seems to be higher which influences all the indicators.
Table 5.6 Results of the simulations of the PBMMR and FBMR
PBMMR FBMR
Methane conversion (%) 98.5 89.7
H2/CO ratio 0.26 1.14
CO selectivity 0.12 0.19
H2 selectivity (%) 96.0 82.0
HRF (%) 93.3 76.0
SF (%) 99.0 92.7
0.00 0.05 0.100%
20%
40%
60%
80%
100%
Met
hane
con
vers
ion
z (m)
FBMR PBMMR
0.00 0.05 0.100%
20%
40%
60%
80%
100%
Nor
mal
ized
hyd
roge
n pe
rmea
tion
z (m)
FBMR PBMMR
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64 Theoretical comparison reactor types for ATR of methane
Based on the isothermal performance, the PBMMR is the better choice for the ATR of methane.
When the adiabatic simulations are considered, a distinct temperature profile is formed in the
PBMMR. Where the PBMMR is performing better than the FBMR in terms of mass transfer, it loses
significantly in heat transfer. The FBMR has excellent heat transfer and also has a uniform
temperature within the bed whereas the PBMMR has a local increase in temperature of over 300 °C.
Heat management is very important for the production of ultrapure hydrogen as the membranes are
not stable at elevated temperatures.
An attempt has been made to apply cooling to the PBMMR but the temperature increase is very local
and it is therefore difficult to tackle this problem. Either a large cooling area should be placed at the
beginning of the bed but this is almost impossible due to the small dimensions (millimeters).
Another possibility is to use two microchannels: one to perform ATR without membranes and a
second one with membranes. The second microchannel will not notice the temperature increase
from the first channel. However, this is not in the scope if this work, where the goal is to have one
reactor performing ATR. So based on the adiabatic performance (with cooling), the FBMR is the
more favored reactor type due to its excellent heat transfer.
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65 Theoretical comparison reactor types for ATR of methane
5.3 Economical
The scaling up of the reactor types to a small industrial scale plant goes impaired with increased
costs. Scaling up will be done to have an equal amount of ultrapure hydrogen produced using the
results of the simulations. The aim is to produce a stream of ultrapure hydrogen of 100 Nm3/h which
is the production of a small scale plant of Hygear.
A preliminary cost analysis has been made and the results can be found in Table 5.7, a detailed
calculation has been added in the Appendix. The preliminary cost analysis is made based on the
isothermal simulations, assuming the heat problems for the PBMMR can be solved. For the FBMR,
30% catalyst loading and a membrane area of 136 m2/m3 showed excellent results compared to the
materials used, so those settings are used. The reactor dimensions were optimized in the simulation
to produce 100 Nm3/h, so no scale-up of the simulation is needed. The PBMMR is modeled as a
microchannel but scale-up in terms of production can be done by multiplying the number of
channels.
The fluidized bed is assessed as a large vessel as can be seen in Figure 5.11 (l). The membranes are
assumed to be placed in the reactor as cylindrical membranes attached to the top of the reactor. With
the used settings, the membranes will use 17% of the reactor area. The PBMMR is assumed to be
square microchannels of 10 cm on two plates separated with a membrane (Figure 5.11 (r)). The
catalyst is assumed to be present as a packed bed in the channel housing and the permeate housing
is empty.
Figure 5.11 Representation of fluidized bed (l) and the microreactor (r)
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66 Theoretical comparison reactor types for ATR of methane
Table 5.7 Preliminary cost analysis for the membrane assisted reactor concepts for a small scale production plant
PBMMR FBMR
From simulation: Flow of ultrapure hydrogen (m3/h) 0.017 103
Channels needed for 100 m3/h 5,900
Wall thickness (cm) 2.54 2.54
Vessel weight (kg) 35,400 450
Reactor costs (€) 285,000 65,000
Amount of catalyst (kg) 25.5 19
Cost of catalyst (€) 25,500 19,000
Effective membrane area (m2) 1.85 3.7
Cost of membranes (€) 18,500 37,000
Total cost estimate (€) 329,000 121,000
The costs of the reactor vessel is done with the method of (Seider et al., 2010). For the FBMR, the
calculation is quite straightforward until the costs of the distributor plate. For the PBMMR several
assumptions have been made to compensate for increased usage of material and the enormous
amount of labor needed to manufacture 5,900 microchannels. The assumptions are mentioned in
the Appendix.
The price of the catalyst per kg is estimated to be € 1000. The catalyst is a 2 wt% Rh/ZrO2. In the
simulations, the kinetics for ATR on a nickel catalyst were used. The activity of the Rhodium catalyst
is seven times higher (Wolbers, 2013) so in practice the catalyst can be diluted or the active
membrane area can be decreased due to the increased hydrogen production. In terms of the stability
of the system, it would be wise to decrease the membrane area.
The price of the membranes is based on the palladium price (Helmi et al., 2014) and the membrane
is assumed to be 4.5 µm in thickness (the active layer thickness of the Tecnalia membranes). The
price of the membranes is estimated to be around € 10,000/m2 of membrane.
From economic point of view, the FBMR is a better option for the ATR of methane. In practice, the
reaction will not be isothermal and cooling has not been taken into account in this preliminary cost
analysis, so the PBMMR will be even more expensive than shown here.
Technische Universiteit Eindhoven University of Technology
67 Theoretical comparison reactor types for ATR of methane
5.4 Conclusion
The PBMMR has been modeled and the steam to carbon ratio has been varied to evaluate its effect.
A higher SCR results in lower methane concentration, thus decreasing the rate of reaction. A lower
SCR is therefore favorable. Adiabatic simulations have been run and have shown a disturbing local
temperature increase. Several cooling methods were applied but no solution has been found.
The FBMR has been modeled and the catalyst loading and specific membrane area have been varied.
It can be concluded that both parameters are important for the overall performance. The more the
better but taking into account the impact on the costs and the stability of the membranes, an
optimum has to be found which for now is chosen to be a catalyst loading of 30% and a specific
membrane area of 136 m2/m3 as the relative increase in methane conversion and HRF is quite small.
Both reactor concepts have their advantages and disadvantages. Based on the simulations, the results
in terms of conversion and hydrogen recovery factor are in favor of the PBMMR concept. The
PBMMR suffers less from mass transfer limitations but is suffering a lot from heat transfer
limitations. The stability of the system is heavily dependent on the temperature and is more
important than high conversions so for that reason, the FBMR is the favored reactor concept.
The preliminary cost analysis showed the FBMR is the better option. Although its less efficient than
the PBMMR, the manufacture of the fluidized bed is significantly easier and therefore the
purchasing costs are much lower.
Technische Universiteit Eindhoven University of Technology
68 Conclusion
6. Conclusion
Several aspects of the membrane assisted reactor for autothermal reforming of methane are
assessed. Hydrogen selective membranes were researched and the permeability constants of two
membranes, provided by ENEA and Tecnalia, have been determined and compared to known Pd-
based hydrogen selective membranes from literature. It can be concluded the membranes perform
better in terms of hydrogen permeation and membrane stability.
A novel catalyst was developed and it is concluded that the Ru/CeZrO2 is mechanically stable and
can be used in a fluidized bed. The stability tests performed showed a decrease in activity of the
catalyst in terms of methane conversion for the autothermal reforming of methane.
Integration of membrane and catalyst has been done with two reactor concepts, the packed bed
membrane microreactor and the fluidized bed membrane reactor. The PBMMR showed low
conversions (< 5%) whereas the FBMR showed conversions up to 90%. The FBMR system with the
Tecnalia membranes showed to be unstable due to leakage of the membrane sealings. The leakage
was found to be caused by the sealings of the membranes which start leaking at temperatures above
500 °C. In fluidization conditions, the increase in leakage is significantly higher. Due to the low
conversions of the microreactor and the leakage of the sealings in the FBMR, the production of
ultrapure hydrogen on lab scale have been unsuccessful.
The reactor types were also compared theoretically. It is concluded the FBMR concept is better for
the production of ultrapure hydrogen through the autothermal reforming of methane due to its
better heat transfer within the bed and the lower cost for purchasing a small scale reactor capable of
producing 100 m3/h of ultrapure hydrogen.
Both the experimental and the theoretical comparison show the FBMR concept is favored for the
autothermal reforming of methane over the PBMMR concept. The favored catalyst choice is the
Ru/CeZrO2 catalyst over the Rh/ZrO2 catalyst due to its lower price and it has proven to be
performing excellent in the FBMR (due to the high amount of catalyst used). The Tecnalia
membranes are favored over the ENEA membranes due to its better hydrogen permeance, although
stability issues of the sealing should be solved.
Technische Universiteit Eindhoven University of Technology
69 Recommendations
7. Recommendations
The sealing of the membrane has room for improvement. For producing ultrapure hydrogen,
selectivity and therefore leakage are very important. The sealings used in this work are tightened in
the radial direction of the membrane. If the force applied is too high, the membrane will break. A
method exists where the force applied is in the longitudinal direction which could avoid tube
cracking (Smart et al., 2012). Smart et al. also reported using the sealing successfully at temperatures
up to 600 °C and pressures up to 6 bar. The sealing is able to be welded to stainless steel fittings
making it viable for use in a fluidized bed. The sealing is patented under US20030146625 (Rusting
et al., 2003). Looking at the direction of the force applied could help improve the sealing, more
pressure can be exerted on the tubular membrane improving the compression of the graphite and
thus the sealing.
The Ru/CeZrO2 catalyst is losing its activity after a few hours of reaction. The morphology of the
catalyst has not been investigated as it was not the scope of this project. Nonetheless, it’s interesting
to investigate if morphology changes are the cause of the decrease in catalytic activity. It could also
help in designing a new catalyst if the cause of deactivation is known. Furthermore it’s also important
investigate the influence of temperature. If deactivation is caused by hotspots, the catalyst would be
suitable for use in a fluidized bed due to the excellent heat transfer.
During the tests with the FBMR, a large difference was observed in the hydrogen permeation
compared to the permeation tests in an empty tube which is most probably caused by concentration
polarization. Unfortunately there is not much known about concentration polarization and its effect
on the overall hydrogen permeation. This is a topic which could help improve the design of the
membrane assisted fluidized bed reactor.
From experimental and theoretical point of view, the microreactor is less promising for the
autothermal reforming of methane. Heat transfer is essential for autothermal reforming, which is
important for the reaction equilibrium and possibly also for the membrane and catalyst.
Experimental tests with the microreactor showed unexpected results with very low conversions. It is
therefore recommended to focus solely on optimizing the fluidized bed membrane reactor as this
concept shows more potential.
Technische Universiteit Eindhoven University of Technology
70 Acknowledgement
Acknowledgement
I would like to thank a few persons for making this thesis possible. First of all Martin van Sint
Annaland for giving me the opportunity to work on this project and for the open-minded attitude
towards me and the project.
For the first seven months my direct supervisor was Lucia Marra. I want to thank her for her
involvement and commitment to my project. Although she could be intense in her communication,
I think I needed that and I enjoyed working with her.
After the first seven months my direct supervisor became Fausto Gallucci. Especially on the
experimental part we’ve had quite some discussions about how to do things and his input was always
well appreciated.
I would like to thank Edwin Zondervan for taking the time to be my external committee member.
I’ve spent quite some time on the lab and had a lot of help with that, especially from the technical
staff. I want to thank Joris Garenfeld for his commitment and help with the setups I used (and
broke). I broke a lot of membranes which caused both you and me quite some headaches but in the
end we could achieve some results. Next to Joris, I would also like to thank Joost Kors and Lee
McAlpine for their technical support.
Next to the technicians, I’ve also had help from Arash Helmi and Kai Coenen with perfoming my
experiments. Not all experiments had the expected outcome and Arash and Kai provided me with
feedback which has been well appreciated. Next to that, Kai has also helped me with this thesis.
Thanks for that.
Special thanks to Ekain Fernandez from Tecnalia for providing expertise on the membranes and the
results on the analysis of the membranes post experiments.
In total I’ve spent around fifteen months at the SMR group. I really enjoyed my time in the group,
maybe a bit too much at some points. I want to thank all the people in the SMR group for making
me feel welcome, for the nice conversations during coffee breaks and borrels and for the pleasant
time.
Technische Universiteit Eindhoven University of Technology
71 Biblography
Biblography
Arzamendi, G., Die, P.M., Gandı, L.M., 2013. Renewable Hydrogen Energy : An Overview 1–17. doi:10.1016/B978-0-444-56352-1.00001-5
Baker, C.G.J., Geldart, D., 1978. An investigation into the slugging characteristics of large particles. Powder Technol. 19, 177–187. doi:10.1016/0032-5910(78)80026-6
Balat, M., Balat, M., 2009. Political, economic and environmental impacts of biomass-based hydrogen. Int. J. Hydrogen Energy 34, 3589–3603. doi:10.1016/j.ijhydene.2009.02.067
Bartels, J.R., Pate, M.B., Olson, N.K., 2010. An economic survey of hydrogen production from conventional and alternative energy sources. Int. J. Hydrogen Energy 35, 8371–8384. doi:10.1016/j.ijhydene.2010.04.035
Chen, W., Hu, X., Wang, R., Huang, Y., 2010. On the assembling of Pd/ceramic composite membranes for hydrogen separation. Sep. Purif. Technol. 72, 92–97. doi:10.1016/j.seppur.2010.01.010
CHEN, Z., GRACE, J., JIMLIM, C., LI, A., 2007. Experimental studies of pure hydrogen production in a commercialized fluidized-bed membrane reactor with SMR and ATR catalysts. Int. J. Hydrogen Energy 32, 2359–2366. doi:10.1016/j.ijhydene.2007.02.036
Deshmukh, S.A.R.K., Heinrich, S., Mörl, L., van Sint Annaland, M., Kuipers, J.A.M., 2007. Membrane assisted fluidized bed reactors: Potentials and hurdles. Chem. Eng. Sci. 62, 416–436. doi:10.1016/j.ces.2006.08.062
Gallucci, F., Chiaravalloti, F., Tosti, S., Drioli, E., Basile, A., 2007. The effect of mixture gas on hydrogen permeation through a palladium membrane: Experimental study and theoretical approach. Int. J. Hydrogen Energy 32, 1837–1845. doi:10.1016/j.ijhydene.2006.09.034
Gallucci, F., Fernandez, E., Corengia, P., van Sint Annaland, M., 2013. Recent advances on membranes and membrane reactors for hydrogen production. Chem. Eng. Sci. 92, 40–66. doi:10.1016/j.ces.2013.01.008
Geldart, D., 1973. Types of gas fluidization. Powder Technol. 7, 285–292. doi:10.1016/0032-5910(73)80037-3
Geldart, D., Abrahamsen, A.R., 1978. Homogeneous fluidization of fine powders using various gases and pressures. Powder Technol. 19, 133–136. doi:10.1016/0032-5910(78)80084-9
Helmi, A., Gallucci, F., van Sint Annaland, M., 2014. Resource scarcity in palladium membrane applications for carbon capture in integrated gasification combined cycle units. Int. J. Hydrogen Energy 39, 10498–10506. doi:10.1016/j.ijhydene.2014.05.009
Hoffmann, A.C., Romp, E.J., 1991. Segregation in a fluidised powder of a continuous size distribution. Powder Technol. 66, 119–126. doi:10.1016/0032-5910(91)80093-X
Holladay, J.D., Hu, J., King, D.L., Wang, Y., 2009. An overview of hydrogen production technologies. Catal. Today 139, 244–260. doi:10.1016/j.cattod.2008.08.039
Kothari, R., Buddhi, D., Sawhney, R.L., 2008. Comparison of environmental and economic aspects of various hydrogen production methods. Renew. Sustain. Energy Rev. 12, 553–563. doi:10.1016/j.rser.2006.07.012
Technische Universiteit Eindhoven University of Technology
72 Biblography
Kunii, D., Levenspiel, O., 1991. Fluidization Engineering, second edi. ed.
Kusakabe, K., Sotowa, K.-I., Eda, T., Iwamoto, Y., 2004. Methane steam reforming over Ce–ZrO2-supported noble metal catalysts at low temperature. Fuel Process. Technol. 86, 319–326. doi:10.1016/j.fuproc.2004.05.003
Liguras, D.K., Kondarides, D.I., Verykios, X.E., 2003. Production of hydrogen for fuel cells by steam reforming of ethanol over supported noble metal catalysts. Appl. Catal. B Environ. 43, 345–354. doi:10.1016/S0926-3373(02)00327-2
Mahecha-Botero, A., Boyd, T., Gulamhusein, A., Comyn, N., Lim, C.J., Grace, J.R., Shirasaki, Y., Yasuda, I., 2008. Pure hydrogen generation in a fluidized-bed membrane reactor: Experimental findings. Chem. Eng. Sci. 63, 2752–2762. doi:10.1016/j.ces.2008.02.032
Marra, L., Wolbers, P.F., Gallucci, F., Annaland, M.V.S., 2013. Development of a RhZrO2 catalyst for low temperature autothermal reforming of methane in membrane reactors. Catal. Today. doi:10.1016/j.cattod.2013.10.069
Mears, D.E., 1971. Tests for Transport Limitations in Experimental Catalytic Reactors. Ind. Eng. Chem. Process Des. Dev. 10, 541–547. doi:10.1021/i260040a020
Mejdell, a. L., Jøndahl, M., Peters, T. a., Bredesen, R., Venvik, H.J., 2009. Effects of CO and CO2 on hydrogen permeation through a ∼3μm Pd/Ag 23wt.% membrane employed in a microchannel membrane configuration. Sep. Purif. Technol. 68, 178–184. doi:10.1016/j.seppur.2009.04.025
Mejdell, A.L., Jøndahl, M., Peters, T.A., Bredesen, R., Venvik, H.J., 2009a. Experimental investigation of a microchannel membrane configuration with a 1.4μm Pd/Ag23wt.% membrane—Effects of flow and pressure. J. Memb. Sci. 327, 6–10. doi:10.1016/j.memsci.2008.11.028
Mejdell, A.L., Peters, T.A., Stange, M., Venvik, H.J., Bredesen, R., 2009b. Performance and application of thin Pd-alloy hydrogen separation membranes in different configurations. J. Taiwan Inst. Chem. Eng. 40, 253–259. doi:10.1016/j.jtice.2008.12.013
Miguel, C.V., Mendes, a., Tosti, S., Madeira, L.M., 2012. Effect of CO and CO2 on H2 permeation through finger-like Pd–Ag membranes. Int. J. Hydrogen Energy 37, 12680–12687. doi:10.1016/j.ijhydene.2012.05.131
Morreale, B.D., Ciocco, M. V., Enick, R.M., Morsi, B.I., Howard, B.H., Cugini, A. V., Rothenberger, K.S., 2003. The permeability of hydrogen in bulk palladium at elevated temperatures and pressures. J. Memb. Sci. 212, 87–97. doi:10.1016/S0376-7388(02)00456-8
Muellerlanger, F., Tzimas, E., Kaltschmitt, M., Peteves, S., 2007. Techno-economic assessment of hydrogen production processes for the hydrogen economy for the short and medium term. Int. J. Hydrogen Energy 32, 3797–3810. doi:10.1016/j.ijhydene.2007.05.027
Mulet, A., Corripio, A.B., Evans, L.B., 1981a. Estimate Costs of Pressure Vessels via Correlations. Chem.Eng. 88, 145–150.
Mulet, A., Corripio, A.B., Evans, L.B., 1981b. Estimate Costs of Distillation and Absorption Towers via Correlations. Chem.Eng. 88, 77–82.
Numaguchi, T., Kikuchi, K., 1988. Intrinsic kinetics and design simulation in a complex reaction network; steam-methane reforming. Chem. Eng. Sci. 43, 2295–2301. doi:10.1016/0009-2509(88)87118-5
Technische Universiteit Eindhoven University of Technology
73 Biblography
Okazaki, J., Ikeda, T., Pacheco Tanaka, D. a, Llosa Tanco, M. a, Wakui, Y., Sato, K., Mizukami, F., Suzuki, T.M., 2009. Importance of the support material in thin palladium composite membranes for steady hydrogen permeation at elevated temperatures. Phys. Chem. Chem. Phys. 11, 8632–8. doi:10.1039/b909401f
Patil, C.S., 2005. Membrane Reactor Technology for Ultrapure Hydrogen Production. Universiteit Twente.
Poling, B., Prausnitz, J., Paul, O.J., Reid, R., 2001. The properties of gases and liquids, 5th ed. McGraw-Hill. doi:10.1036/0070116822
Rusting, F., Pex, P., Peters, J., 2003. Sealing socket and method for arranging a sealing socket to a tube. US20030146625.
Sandler, H.J., Luckiewicz, E.T., 1987. Practical process engineering: a working approach to plant design, McGraw-Hill books in process engineering. McGraw-Hill B. Co.
Seider, W.D., Seader, J.D., Lewin, D.R., Widagdo, S., 2010. Product and Process Design Principles: Synthesis, Analysis, and Evaluation, 3rd ed, Product and process design principles: synthesis, analysis, and evaluation. Wiley.
Smart, S., Vente, J.F., Diniz da Costa, J.C., 2012. High temperature H2/CO2 separation using cobalt oxide silica membranes. Int. J. Hydrogen Energy 37, 12700–12707. doi:10.1016/j.ijhydene.2012.06.031
Stainless Steel - Grade 316L - Properties, Fabrication and Applications [WWW Document], 2013. URL http://www.azom.com/article.aspx?ArticleID=2382 (accessed 6.1.14).
Tanimoto, H., Chiba, S., Chiba, T., Kobayashi, H., 1981. Jetsam Descent Induced by a Single Bubble Passage in Three-dimensional Gas-Fluidized Beds. J. Chem. Eng. Japan 14, 273–276.
Trimm, D.L., Lam, C.-W., 1980. The combustion of methane on platinum—alumina fibre catalysts—I. Chem. Eng. Sci. 35, 1405–1413. doi:10.1016/0009-2509(80)85134-7
Vadrucci, M., Borgognoni, F., Moriani, A., Santucci, A., Tosti, S., 2013. Hydrogen permeation through Pd–Ag membranes: Surface effects and Sieverts’ lawVadrucci, M., Borgognoni, F., Moriani, A., Santucci, A., & Tosti, S. (2013). Hydrogen permeation through Pd–Ag membranes: Surface effects and Sieverts' law. International Journa. Int. J. Hydrogen Energy 38, 4144–4152. doi:10.1016/j.ijhydene.2013.01.091
Wolbers, P.F., 2013. Development of an auto thermal reactor for the production of hydrogen. Eindhoven University of Technology.
www.reforcell.eu [WWW Document], 2013. URL www.reforcell.eu (accessed 10.18.13).
Technische Universiteit Eindhoven University of Technology
74 Appendix
Appendix
A. Comparison of umf of CeZrO2 with literature
The experimentally determined minimal fluidization velocity has been compared to literature to
easily determine the required gas velocities in the fluidized bed. To determine which correlation
described the minimal fluidization best, parity plots were made:
Figure 1 parity plots comparing the experimentally determined minimal fluidization velocity with the predicted values according
to four different correlations
The coefficient of determination for the parity plots was determined and the correlation of Wen and
Yu proves to be describing the minimal fluidization velocity best.
Table 1 Coefficient of determination (R-squared) for the observed data and predicted model values versus the fitted line y=x
Particle fraction
(µm)
Thonglimp (1981) Richardson (1971) Wen and Yu
(1966)
Grace (1982)
90-125 0.812 0.800 0.562 0.699
125-180 0.770 0.550 0.900 0.310
180-250 0.631 0.320 0.916 -0.003
250-315 0.451 -0.080 0.880 -0.638
315-355 -0.311 -1.467 0.667 -2.655
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75 Appendix
B. Parameters used in the packed bed reactor tests
Table 2 Parameters as used in the stability test for ATR of methane in a packed bed reactor
Parameter Value
Catalyst (bulk) density cat [kg/m3] 3450
Catalyst mass catm [kg] 30·10-6
Total molar feed flow tot [mol/s] 0.07
Molar fraction of methane in the feed 4
yCH [-] 0.23
Methane conversion (experimental) 4CH [-] 0.5
Surface-to-volume ratio ca [m2/m3] 5.3·103
Bed porosity [-] 0.4
Temperature T [K] 923.15
Pressure p [Pa] 1.8·105
dynamic viscosity [Pa·s] 3.71·10-6
Gas density of feed composition (923.15 K) [kg/m3] 0.55
Binary mass diffusivity D [m2/s] 5.48·10-5
Superficial gas velocity v [m/s] 1.76
Bed height L [m] 5·10-3
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76 Appendix
C. Graphs FBMR with Tecnalia membranes
Figure 2 Steam methane reforming (SMR) in a FBMR with five tubular Tecnalia membranes at 500 and 550 °C and a pressure
of 1.3 bar. Feed: SCR of 3 and an NCR of 8.4 with a total flow rate of 10.3 Nl/min.
Figure 3 Steam methane reforming (SMR) in a FBMR with five membranes at 550 and 600 °C. Feed: SCR of 3 and an NCR of
8.4 with a total flow rate of 10.3 Nl/min to ensure fluidization.
The graphs in Figure 2 and Figure 3 look stable. The permeate flow was stable which was also
analyzed. From the analysis of this stream it could be seen that the CO level in the permeate stream
was increasing very slowly (~3 ppm/h at 550 °C). The next day, the same reaction was performed.
The leakage was significantly higher than the day before. The membrane sealings were not exposed
to oxygen so the leakage is also caused by other factors. In Chapter 2.3.2.2 we already saw that the
leakage increased at 600 °C.
Figure 4 Autothermal reforming of methane (ATR) in a FBMR with five membranes at 600 °C. Feed: SCR of 3, OCR of 0.25
and an NCR of 8.2 with a total flow rate of 10.3 Nl/min to ensure fluidization.
0 50 100 150 200 250 3000
2
4
6
8
10
12
14
16
18
CO CO2 CH4 H2
Ret
enta
te c
ompo
sitio
n (v
ol%
)
Time
0
20
40
60
80
100
Reaction and separation550 °C CO in permeate = 70 ppm
CH4 conversion (%)
CH
4 co
nver
sion
(%)
Reaction and separation500 °C CO in permeate = 50 ppm
0 50 100 150 200 250 3000
2
4
6
8
10
12
14
16
18
20
Ret
enta
te c
ompo
sitio
n (v
ol%
)
Time (min)
CO CO2 CH4 H2
0
20
40
60
80
100
Reaction and separation600 °C CO in permeate = 160 to 200 ppm
Reaction and separation550 °C CO in permeate = 120 ppm
CH4 conversion (%)
CH
4 co
nver
sion
(%)
0 50 100 150 200 250 3000
2
4
6
8
10
12
14
16
18
20
Ret
enta
te c
ompo
sitio
n (v
ol%
)
Time (min)
CO CO2 CH4 H2 CH4 conversion (%)
0
20
40
60
80
100
Reaction and separation600 °C CO in permeate = >800 ppm
CH
4 co
nver
sion
(%)
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77 Appendix
D. Preliminary cost analysis
The preliminary cost analysis has been performed on the reactor vessel, catalyst and membranes.
The cost prediction of the reactor vessel have been made using the method mentioned in the book
of Seider (Seider et al., 2010) which gives an estimation based on the material used and the weight
and size of the vessel.
The desired production capacity is 100 m3/h, which corresponds to a small scale hydrogen
production plant.
PBMMR FBMR
From simulation:
Flow of ultrapure hydrogen (m3/h) 0.017 103
For the desired 100 m3/h:
Channels needed 5,900
Wall thickness (cm) 2.54 2.54
Vessel weight (kg) 35,400 450
Reactor costs (€) 285,000 65,000
Amount of catalyst (kg) 25.5 19
Cost of catalyst (€) 25,500 19,000
Effective membrane area (m2) 1.85 3.7
Cost of membranes (€) 18,500 37,000
Total cost estimate (€) 329,000 121,000
For calculating the minimal thickness of the fluidized bed reactor the following formula was used
for calculating the wall thickness to withstand the internal pressure:
2 1.2
d ip
d
P DtSE P
1
In which iD is the internal diameter of the vessel, S the maximum allowable stress of the shell
material at the design temperature, E the fractional weld efficiency and dP the internal design
gauge pressure. The internal design gauge pressure can be calculated using a correlation by Sandler
et al., this correlation calculates a safe design pressure based on the desired operational pressure,
OP is the operating pressure (Sandler and Luckiewicz, 1987):
2exp{0.60608 0.91615[ln( )] 0.0015655[ln( )] }d O OP P P 2
Equation 1 is suitable for calculating the thickness of a horizontal pressure vessel, but does not
account for effects of wind or an earthquake on a vertical vessel. An additional thickness is therefore
added to account for effects on vertical vessels:
Technische Universiteit Eindhoven University of Technology
78 Appendix
2
2
0.22( 18)Ow
O
D LtSD
3
In which L is the vessel height and OD the outside diameter of the vessel. Adding the two calculated
thicknesses gives the average vessel wall thickness vt that should be used:
v p wt t t 4
With the design specs used, 8 bar, the wall thickness is calculated at 0.20 cm, which is mechanically
not viable. If this is the case, (Seider et al., 2010) recommend a wall thickness of 0.65 cm. As
explosive gases are produced and to implement more safety, a thickness of 2.54 cm (1 inch) is chosen.
The weight of the reactor can then be calculated by the following formula, which incorporates
elliptical heads:
( )( 0.8 )i v i vW D t L D t 5
Where is the density of the vessel material which is stainless steel 316L. Stainless steel 316L is
chosen for its properties. It’s corrosion resistant, has a high toughness and offers higher stress to
rupture and tensile strength at elevated temperature. SS316L is chosen over SS316 due to its lower
carbon content. SS316 is more prone to carbide precipitation between 425 and 860 °C whereas
SS316L is more resistant to carbide precipitation in this temperature region due to its lower carbon
content in the steel(“Stainless Steel - Grade 316L - Properties, Fabrication and Applications,” 2013).
The purchase cost of the reactor vessel are calculated using the following formula which are
determined empirically by (Mulet et al., 1981a, 1981b):
P M V PLC F C C 6
Where MF is the materials-of-Construction Factor (2.1 for SS316), VC the costs for the vessel and
PLC the costs for additional platforms and ladders which are given by the following equations:
2exp{7.0132 0.18255[ln( )] 0.02297[ln( )] }VC W W 7
0.73960 0.70684361.8( ) ( )PL iC D L 8
With the given reactor dimensions this results in a purchase cost of € 55,000. This does not include
the cost of the distributor plate. The distributor plate is important for a fluidized bed and needs to
be able to distribute the feed evenly over the area of the bed. It is assumed a distributor plate with a
diameter of 0.42 m including pre-feed installation costs € 10,000.
Methods for calculating the costs for a microreactor are not present. To have an indication, the costs
are calculated on the basis of the total reactor volume (divided over the channels) using the same
method as used for the fluidized bed. Considering a microreactor is much more detailed in
manufacturing, it is assumed the weight per reactor volume is 100 times higher based on the fact
the reactor volume to steel ratio is opposite of a fluidized bed and the labor in manufacturing a
microreactor is significantly higher. Furthermore the packed bed has to be inserted in the
microchannels in such a way the pressure drop in every separate microchannel should be the same.
The catalyst is Rh/ZrO2 and was previously made in-house. The costs for preparing the catalyst is
estimated to be € 1000/kg. The membranes used are Pd-based membranes. The cost of palladium
Technische Universiteit Eindhoven University of Technology
79 Appendix
is € 25,000/kg (Helmi et al., 2014). In the simulations, the membranes by Tecnalia were used which
have an active layer thickness of 4.5 µm. This corresponds to a price of € 1000/m2 of membrane
based on material cost. The assumption is made that the cost will multiply by 10 for preparing the
active layer on a ceramic support and preparing them for use in the desired reactor.