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Autothermal reforming of methane in membrane assisted reactors

Ottenheijm, I.N.

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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.


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


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


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


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


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


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


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.


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.


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].



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).



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



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



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



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)



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)



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.



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)



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



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.



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)



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.


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.



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.



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)



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)



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.



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.



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)



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



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)



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



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.



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



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 [-



], 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



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



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)



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.



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



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



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.


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.



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.



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.



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



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.



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)



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



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.



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).



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.



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


CH

4 co

nver

sion

(%)



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 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)


0

50

100

CH

4 co

nver

sion

(%)



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



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.



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.



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)



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.



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



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.



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.



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.


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).



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



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.



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



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



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



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



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



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.



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)



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.



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.


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.


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.


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.


71 Biblography

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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


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


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

(%)


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


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)


0

20

40

60

80

100

Reaction and separation600 °C CO in permeate = >800 ppm

CH

4 co

nver

sion

(%)


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:


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


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.

autothermal reforming of methane in membrane ... - tu/e · in this work, steam methane reforming is...

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