Novel H2 manufacturing routes. Simulation and thermodynamic analysis of membrane based processes

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Novel H

2

Manufacturing Routes

Simulation and Thermodynamic Analysis of Membrane

Based Processes

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Novel H

2

Manufacturing Routes

Simulation and Thermodynamic Analysis of Membrane

Based Processes

Proefschrift

ter verkrijging van de graad van doctor aan de Technische Universiteit Delft

op gezag van de Rector Magnificus prof. dr. ir. J. T. Fokkema voorzitter van het College voor Promoties

in het openbaar te verdedigen op maandag 27 september om 10:30 uur

door

Peijun JI

Master of Science in Chemical Engineering, Beijing University of Chemical Technology, China

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Prof. dr. ir. J. de Swaan Arons

Samenstelling promotiecommissie:

Rector Magnificus Voorzitter

Prof. dr. ir. J. de Swaan Arons Technische Universiteit Delft, promotor

Prof. dr. S. Kjelstrup Norwegian University of Science and Technology

Dr. J. J. C. Geerlings Shell Global Solutions International B.V.

Prof. dr. ir. M. -O. Coppens Technische Universiteit Delft

Prof. dr. J. Schoonman Technische Universiteit Delft

Prof. dr. ir. J. Grievink Technische Universiteit Delft

Dr. ir. H. van der Kooi Technische Universiteit Delft

ISBN: 9090184929

All rights reserved. No part of the material protected by this copyright notice may be reproduced or utilized in any form or by any means, electronic or mechanical, including photocopy, recording or by any information storage and retrieval system, without permission from the publisher.

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1

Summary I

Summary in Dutch III

Summary in Chinese V

1. Introduction 1

1.1. Producing H2 and its application in fuel cell technology 1

1.2. Membrane reactors 3

1.2.1. The O2-membrane catalytic partial oxidation (CPO) reactor 3

1.2.2. H2-membrane reactors 4

1.2.3. The two-membrane CPO reactor 5

1.3. Objective and approach 6

1.4. Outline of the thesis 8

2. Data generation by simulation for process analysis and optimisation 12

2.1. Introduction 12

2.2. Simulation of the conventional CPO reactor 13

2.2.1. Kinetics 13

2.2.2. Reactor model for the conventional CPO reactor 15

2.2.3. Solution procedure 17

2.2.4. Oxidation state of the catalyst 18

2.2.5. Validity of the conventional CPO reactor model 20

2.3. Simulation of membrane CPO reactors 22

2.3.1. Reactor model for membrane CPO reactors 22

2.3.2. The O2-membrane CPO reactor 26

2.3.3. The H2-membrane CPO reactor 26

2.3.4. The two-membrane CPO reactor 29

2.3.5. Validity of the membrane CPO reactor model 31

2.4. Simulation of conventional WGS reactors 34

2.4.1. Kinetic model 34

2.4.2. Simulation models 35

2.4.3. Validity of the models 35

2.4.4. Simulation of the H2-membrane WGS reactor 37

2.5. Co-current and counter-current mode for a reactor with a H2-membrane 38

2.6. Discussion and conclusion 42

3. Comparison of conventional and membrane CPO reactors 45

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reactor

3.2.2. Effect of pressure 49

3.3. The H2-membrane CPO reactor 51

3.3.1. Fraction of H2 recovered (on CO+ H2 basis)

σ

51

3.3.2. Production rate of useful products 51

3.4. The two-membrane CPO reactor 57

3.4.1. Inlet temperature of air 57

4. Thermodynamic analysis of integrated CPO processes with preset inlet

conditions and methane conversion 63

4.2. Brief introduction of thermodynamic analysis 63

4.2.1. Exergy and thermodynamic analysis 63

4.2.2. Calculating the exergy of a gas mixture 64 4.3. Simulation models for other key units (excluding the CPO reactors and

WGS reactors) 66

4.3.1. The PEM fuel cell 66

4.3.2. The selective oxidation reactor 67

4.3.3. The heat exchanger, mixer and evaporator 67

4.3.4. The compressor 67

4.3.5. The catalytic burner 68

4.4. Process considerations 68

4.4.1. Boundaries and starting points 68

4.4.2. Heat recovery and heat integration 70

4.5. Thermodynamic analysis of integrated processes 76 4.5.1. Definition of net work-output and overall exergy efficiency 76 4.5.2. Effect of the inlet temperature of air 78

4.5.3. Effect of the inlet pressure 82

4.5.4. Comparison of four integrated processes 89

4.6. Exergy depletion in four integrated processes 92

5. Thermodynamic analysis of an integrated CPO process with constrained

geometry 97

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3

5.3.1. Overall exergy efficiency 99

5.3.2. Effect of the inlet rate of methane 101

5.3.3. Effect of molar rate of sweeping gas 105

5.3.4. Effect of inlet ratio CH4 /H2O 107

5.4. Discussion and conclusion 110

6. Brief analysis of the conventional steam reforming (SR) process 111

6.1. Introduction 111

6.2. Process description 111

6.3. Simulation of the steam reformer 113

6.4. Simulation and thermodynamic analysis results 116

6.5. Conclusion 118

7. Summary of results and main conclusions 119

8. Outlook 122

List of symbols 124

Acknowledgment 128

Curriculum vitae 129

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Summary

Nowadays hydrogen is gaining more and more attention mainly because it is generally regarded as an important future fuel. Although H2 can be produced from a wide variety of resources using a range of different technologies, natural gas is generally preferred and will remain in the near future the major feedstock for the manufacture of H2. Catalytic partial oxidation (CPO) is seen as a key technology for the conversion of methane in H2.

The H2 produced by the conventional CPO process is mixed with other gases. For this fuel gas mixture rich in H2, the fuel cell does not convert at least 15% of the H2 that has been fed into it. Membrane reactors can be used to produce pure H2, and then the utilization of the hydrogen produced can be enhanced, in the mean time, the process can be simplified.

The membrane reactors which have been studied, include the O2-membrane CPO reactor, the H2-membrane CPO reactor, the two-membrane CPO reactor (a CPO reactor with O2 and H2 permeable membrane), and the H2-membrane water-gas-shift (WGS) reactor. Three integrated processes around these membrane reactors are proposed and compared with the conventional CPO process based on the results of a thermodynamic analysis.

For performing a thermodynamic analysis, a wide range of data is required. To fulfil the data requirement, the simulation has to be accomplished, especially for the CPO and WGS reactors, without and with membranes. The simulation work is based on the kinetics of the reaction and membrane permeation mechanisms. With the simulation models the profiles of temperature and molar fraction of different species along the reactor’s axial coordinate have been simulated. The validity of the simulation work has been confirmed by the comparison of simulation results with the experimental data.

Because of the key role played by the CPO reactors in each corresponding integrated processes, the production rates of useful products by different CPO reactors have been simulated and compared. The effect of each adjustable inlet condition on the output of the CPO reactors has been shown and discussed.

On the basis of the simulation work and the knowledge of the ‘characteristics’ of the conventional CPO reactor and the membrane CPO reactors, the thermodynamic analysis of the integrated conventional and membrane CPO methane conversion processes has been carried out. The results have been used to

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process has been identified. The resulting exergy balance for each process has been used to determine where and to what extent the exergy, that part of the energy contained that is available for useful work (such as electricity from a fuel cell), is dissipated.

Also on the basis of the simulation work, the thermodynamic analysis of an integrated H2 production process with a fixed-geometry H2-membrane reactor has been described. The effect of different inlet conditions on the overall exergy efficiency of the process has been shown and discussed. The simulation and thermodynamic analysis have provided a quantitative tool to get insight into and to understand the various hydrogen production processes.

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Samenvatting

Tegenwoordig krijgt waterstof meer en meer aandacht hoofdzakelijk omdat het in het algemeen als belangrijke toekomstige brandstof wordt beschouwd. Hoewel H2 uit een grote verscheidenheid van grondstoffen kan worden geproduceerd gebruikmakend van een veelheid van verschillende technologieën, heeft het gebruik van aardgas in de nabije toekomst de voorkeur als de belangrijkste grondstof voor de vervaardiging van H2. Katalytische partiële oxidatie (CPO) wordt gezien als de belangrijkste technologie voor de omzetting van methaan in waterstof.

H2, geproduceerd via het conventionele CPO proces, is gemengd met andere gassen. Voor dit brandstofmengsel, rijk aan H2, is de omzetting van H2 in de brandstofcel ten hoogste 85%. Membraanreactoren kunnen worden gebruikt om zuivere H2 te produceren terwijl meer van de geproduceerde waterstof kan worden benut en het proces kan worden vereenvoudigd.

De membraanreactoren, die zijn bestudeerd, zijn de O2-membraan CPO reactor, de H2-membraan CPO reactor, de twee-membraan CPO reactor (CPO reactor met een O2 en een H2 membraan), en een H2-Membraan water-gas-shift (WGS) reactor. Drie rond deze membraanreactoren geïntegreerde processen worden voorgesteld en vergeleken met het conventionele CPO proces op basis van de resultaten van een thermodynamische analyse.

Voor het uitvoeren van een thermodynamische analyse zijn veel gegevens vereist. Om in de behoefte aan gegevens te voorzien moeten deze gesimuleerd worden, vooral voor de CPO en WGS reactoren met of zonder membranen. Het simulatiewerk is gebaseerd op de kinetiek van de reacties en op de permeatiemechanismen van de membranen en het reactor model. Met de simulatiemodellen zijn de temperatuur- en de molfractie-profielen van de verschillende stoffen langs de asrichting van de reactoren gesimuleerd. De geldigheid van het simulatiewerk is bewezen door vergelijking van de simulatieresultaten met de experimentele gegevens.

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productiesnelheden van de diverse waardevolle producten van de verschillende CPO reactoren gesimuleerd en vergeleken. Het effect van elke regelbare inlaatconditie op de output van de CPO reactoren is aangetoond en besproken. Op basis van het simulatiewerk en de kennis van de ' karakteristieken ' van de conventionele CPO reactor en de membraan CPO reactoren, is de thermodynamische analyse van de geïntegreerde conventionele en membraan CPO processen voor de omzetting van methaan uitgevoerd. De resultaten zijn gebruikt om de hele processen te ontwerpen en te optimaliseren rond elke CPO reactor. Het meest efficiënte proces is geïdentificeerd. De exergie balans is voor elk proces gebruikt om te bepalen waar en in welke mate exergie, dit is het deel van de energie dat maximaal als arbeid beschikbaar is (zoals elektriciteit van een waterstof gevoede brandstofcel), verloren is gegaan.

Op basis van het simulatiewerk is ook de thermodynamische analyse van een geïntegreerd H2 productieproces met een H2-membraan reactor met een vastgelegde geometrie beschreven. Het effect van verschillende inlaatcondities op de exergie efficientie van het hele proces is bepaald en besproken. De simulatie en de thermodynamische analyse waren een kwantitatief hulpmiddel om inzicht te krijgen in verschillende waterstof productieprocessen en om deze te begrijpen.

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氢作为重要的未来能源载体受到越来越广泛的重视。尽管氢气可以由多种原料通过多种途径来制取，但是天然气被广泛认为是现在及将来用于制取氢气的主要原料。催化部份氧化(Catalytic partial oxidation,CPO)被认为是转化天然气制氢的一种重要技术。在传统的催化部份氧化过程中，所制取的氢气与其它气体相混合。对于这种富含氢气的混合气，在氢电池中，至少15%的输入氢气不能转化为电能。解决这个问题的一个办法是应用膜反应器来制备纯氢，从而可以提高氢气利用率并简化生产工艺过程。在本工作中，所研究的膜反应器包括：可透过氧的CPO膜反应器，可透过氢的CP O膜反应器，可透过氧及可透过氢的CPO双膜反应器，可透过氢的Water-gas-shift(WGS)反应器。围绕这些膜反应器，本工作提出了三个新的制氢工艺过程。通过对反应器的模拟以及对生产工艺过程的热力学分析，对这三个新过程及传统过程进行了比较。对传统反应器和膜反应器的模拟数据为生产工艺过程的热力学分析提供了坚实的基础。反应器模拟是基于反应动力学机理及膜透过机理。模拟结果包括沿反应器轴向的温度分布及不同组份的摩尔分率分布。通过对模拟结果和实验结果的比较，所用模拟方法的可靠性得到了验证。因为CPO反应器的重要性，所以本工作首先对不同CPO反应器的有用产品产率进行了分析和比较。模拟结果还包括可变进口条件(如压力和反应物料进口流率比等) 对CPO反应器有用产品产率的影响。基于对各CPO反应器模拟结果及对反应器特性的认识，对传统CPO过程及新过程进行了热力学分析，其结果用于生产工艺过程的设计和优化以及对过程的评价。对固定几何尺寸的可透过氢的CPO膜反应器进行了模拟，并对含有这样反应器的完整工艺过程进行了热力学分析。同时还分析和讨论了不同进口条件对制氢速率及过程热力学效率的影响。模拟和热力学分析结果对于更深入地认识和理解不同的制氢过程提供了量化工具。

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1. Introduction

Hydrogen is primarily used as feedstock, intermediate chemical, or specialty chemical. Nowadays hydrogen is gaining more and more attention mainly because it is generally regarded as an important future fuel. Widespread use of hydrogen as an energy source in this 21st_{century could help alleviate concerns about energy} security, global climate change, and air quality [1,2,3]_{. Through fuel cells, hydrogen} has the potential to power everything from a single family home to a whole neighbourhood to an entire state or region.

1.1. Producing H

2

and its application in fuel cell technology

The quantity of natural gas reserves is huge and approximates the energy equivalent of the known crude oil reserves. In the past decade, global research activities have been unfolded both in industry and academia to develop economic processes for the conversion of large natural gas reserves to valuable products, such as methanol and hydrogen.

Although H2 can be produced from a wide variety of resources using a range of different technologies, such as thermal or biological conversion of biomass, splitting water using electricity, heat or light, and gasification of fossil fuels, like coal, natural gas is generally preferred and will remain in the near future the major feedstock for the manufacture of H2 due to its availability, because it is clean, and low-cost for transportation [4,5,6]_.

There are also significant environmental drivers for the conversion of natural gas. Current practices of venting or flaring associated with natural gas contribute to global warming. Natural gas conversion to high-grade, high-purity fuels affords an opportunity to reduce combustion pollutants [7]_.

The first step in natural gas (assumed pure methane, CH4) conversion is often the production of syngas, a mixture of hydrogen, carbon monoxide and carbon dioxide. Syngas is mainly produced by steam reforming (SR, Eq.(1.1))[8]_.

H

4 2 3

CH +H OUCO+ ₂ 0

1 206.2 /

∆H = kJ mol (1.1) Since the steam reforming reaction is strongly endothermic, a steam reformer normally includes a gas-fired furnace and parallel reactor tubes [9]_.

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_________________________________________________________________________________ An alternative technique for producing syngas is catalytic partial oxidation (CPO). In a CPO reactor, methane is converted into syngas by reacting with steam and the oxygen of air. This step is also called auto-thermal reformer, which means the heat required by the endothermic steam reforming is supplied by the exothermic methane combustion reaction (Eq.1.2) within the reactor, so that no parallel furnace is needed. In the CPO and SR reactors, there is also the water-gas shift (WGS) reaction (Eq.1.3) taking place.

4 2 2 2 2 CH + O UCO + H₂O 0 2 802.3 / ∆H = − kJ mol (1.2) 2 CO H O+ UCO₂+H2 ∆H30 = −41.2 kJ mol/ (1.3)

As illustrated in Fig.1.1, in order to convert most of the CO in the syngas into H2, a high temperature WGS converter (HWGS) and a low temperature WGS converter (LWGS) are placed after the CPO or SR reactor. The operating temperatures of the two WGS converters are around 720 K and 480 K, respectively [10]_{. Water is injected into the mixers to cool down the hot gas streams before they} enter the WGS converters. Water addition increases the steam to CO ratio that drives the WGS reaction towards a higher CO conversion.

CH4 H2O SR HWGS LWGS SOR H₂ Fuel cell CO2 Electricity

a

CH4 H2O Air

CPO HWGS LWGS SOR H2 Fuel cell

CO2 N2

Electricity

b

Fig.1.1. The conventional CH4 conversion processes producing H2 as feedstock for a fuel

cell.

Among the several different types of fuel cells, the polymer electrolyte membrane (PEM) fuel cell is the most promising option due to its lower operating temperature, higher power density, higher energy efficiency and fast response [2,3]_. CO is a poison for the anode catalyst of the PEM fuel cell. The catalyst within the fuel cell can tolerate a maximum of 40 ppm CO, which cannot be achieved by the

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_________________________________________________________________________________ WGS converters only. Hence, a selective oxidation reactor (SOR) is normally placed after the LWGS reactor to further reduce the CO concentration within the H2 rich stream. In the selective oxidation reactor, a small amount of air is added to the stream and the CO is selectively oxidized in a fixed platinum catalyst bed while consuming a minimum amount of H2 [11].

As the PEM fuel cell normally operates in the temperature range of 353~363 K, the fuel gas out of the SOR should be cooled down to that temperature before going into the fuel cell.

f

The H2 produced by the SR or CPO process is mixed with other gases, like CO2, N2, steam and unconverted CH4. For this fuel gas mixture rich in H2, the PEM fuel cell does not convert some 25% of the H2[12,13] that is fed into it. In this example, the percentage of H2 converted in the fuel cell

η

, which is defined as the percentage of the converted H2 to all the H2 fed into the fuel cell, is 75%. By increasing the equipment investment in the fuel cell, the percentage of H2 converted in the fuel cell

η

_f of the H2 rich gas mixture can be increased up to 85%[14]_.

If H2 is separated from the gas mixture, this problem can be avoided [15,16]. One of the solutions is the use of membrane to separate H2.

1.2. Membrane reactors

Membrane reactors continue to be of interest for reaction engineers, due to their ability to carry out simultaneously reaction and separation [17]_{. Such reactors can} achieve preferential removal of products, or can regulate the stoichiometric feed of reactants. The development of O2-membrane and H2-membrane technologies provides the possibility to improve the methane conversion processes.

1.2.1. The O2-membrane CPO reactor

O2 permeation fluxes considered to be economically feasible have been found through membranes derived from the perovskite structure ABO3 or the brownmillerite structure A2B2O5, where A is preferably a rare earth or alkaline-earth (e.g. La , Sr, Ba, Ca) and B a transient-metal (e.g. Ce, Co, Fe). The membrane performance is strongly dependent on the perovskite stoichiometry, which governs either the oxygen transfer rate or the membrane stability. The potential offered by new perovskites showing comparatively higher permeability has still to be fully explored [18-20]_.

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_________________________________________________________________________________ Air Air CH4 H2O O2- membrane Catalyst Syngas Exhaust air Exhaust air

Fig.1.2. Configuration of an O2-membrane CPO reactor.

The most practical application of the O2-membrane is the O2-membrane CPO reactor [21]_{, as shown in Fig.1.2. In an O}

2-membrane CPO reactor, air and a methane/water mixture are flowing through the two separated compartments. The oxygen from air, diffusing through the membrane along the reactor’s axial coordinate, reacts with methane to form syngas. Various advantages of an O2 -membrane CPO reactor are: (1) employing air as oxidant but eliminating N2 contamination in the product; (2) improving safety management since the ceramic membrane avoids premixing of oxygen and hydrocarbons before the occurrence of the reaction; (3) controlling the supply of oxygen along the reactor’s axial coordinate, mitigating the formation of ‘hot spots’ [22]_{; (4) eliminating NO}

x emission since atmospheric nitrogen is excluded from permeating through the membrane. Therefore, O2-membrane reactors have been paid considerable attention in recent years [23-26]_.

1.2.2. H2 -membrane reactors

Membranes made from palladium-rich alloys such as Pd-Ag have been used in industry to separate H2 from H2-containing mixtures or to make ultra pure hydrogen. Many investigators have studied H2-membrane CPO and H2-membrane WGS reactors [27-30]_{. The configurations of a H}

2-membrane CPO reactor and a H2 -membrane WGS reactor are shown in Figs.1.3a and 1.3.b, respectively.

The H2-membrane reactors combine the separator and the reactor in one process step. By selectively removing the H2 through the membrane, methane conversion or carbon monoxide conversion in the H2-membrane reactors can be increased compared with the conversion of a conventional CPO reactor or a conventional WGS reactor, respectively. Directly producing high purity hydrogen is one of the main advantages of the H2-membrane reactor.

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_________________________________________________________________________________

CH4

H2O

Air

Sweeping gas Sweeping gas + H2

H2 -membrane Catalyst

Rejected fuel gas

a. H2 -membrane CPO reactor

Syngas H2O

H2 -membrane

Catalyst Rejected fuel gas

b. H2 -membrane WGS reactor

Fig.1.3. Configurations of H2-membrane reactors (the sweeping gas in a co-current mode).

1.2.3. The two-membrane CPO reactor

As shown in Fig.1.4, the two-membrane CPO reactor integrates various

processing steps in a single reactor. The reactor is divided into three compartments. Like in the O2-membrane CPO reactor, in the two-membrane CPO reactor, O2 is transported across the O2-membrane to the reaction side, where it reacts with methane. The formed syngas is also exposed to the H2-membrane. The hydrogen is transported to the permeate side of the H2-membrane by the difference in H2 partial pressure. Due to the removal of H2 from the reaction zone, more H2 is formed by the reforming and shift reactions.

The two-membrane CPO reactor combines the advantages of the O2-membrane CPO reactor and that of the H2-membrane CPO reactor. Minish, et al [31] investigated the techno-economic feasibility of the two-membrane system. Based

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_________________________________________________________________________________ on initial feasibility assessment, they found that there are no technical issues that will prevent the development of the integrated two-membrane reactor. Chen, et al [32]_{simulated a novel circulating fast fluidized-bed reactor with O}

2 and H2 -membranes. By optimizing the number of H2-membranes, the number of O2 -membranes, the oxygen feed rate and the steam-to-carbon ratio, they found that the hydrogen productivity in the novel reactor is about 8 times higher than that in typical industrial fixed-bed steam reformers.

Air Sweeping gas CH4 H2O O2 membrane H2 membrane Catalyst Exhaust air Sweeping gas + H 2

Rejected fuel gas

Fig.1.4. Configuration of two-membrane CPO reactor (the sweeping gas in a co-current mode).

1.3. Objective and approach

Many researchers have studied membrane CPO reactors for methane conversion due to their advantages over conventional CPO reactors. However, so far studies of integrated processes using these membrane reactors producing H2 have not been reported.

The objective of this work is to compare various integrated CPO processes including a conventional process and three possible membrane processes through a thermodynamic analysis. The thermodynamic analysis of the processes, per step and in overall, helps to determine where and to what extent energy is dissipated because of irreversibilities [33-35]_{. Therefore, based on the thermodynamic analysis} results on the inefficiency of the units in each process, process improvement can be proposed, and the most efficient process can be identified.

Three possible membrane processes producing H2 for fuel cell application are shown in Fig.1.5. For the membrane processes shown in Fig 1.5a and 1.5c, an O2 -membrane is applied in the CPO reactor, due to the advantages of an O2-membrane reactor described in section 1.2.1. For all three processes, a H2-membrane is applied either in the CPO reactor or the WGS reactor to separate H2. Steam is chosen as the sweeping gas in the permeate side of the H2-membrane reactors. The

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_________________________________________________________________________________ presence of steam does not affect the percentage of H2 converted, because water can move across the membrane in the fuel cell [36]. The separated H2 can be converted to depletion in the fuel cell.

O2-Membrane CPO H2-Membrane WGS Sweeping gas H2 Air Two-membrane CPO H2-membrane CPO Air Fuel cell Fuel cell CH4 H2O CH4 H2O Air CH4 H2O Fuel cell a b c H2 -membrane O2 -membrane Exhaust air Syngas Sweeping gas Electricity

Rejected fuel gas

Furnace

Rejected fuel gas

Furnace

Rejected fuel gas

Electricity Furnace Electricity Exhaust air Sweeping gas Sweeping gas H₂

Sweeping gas Sweeping gas_H 2

Fig.1.5. Scheme of three possible membrane processes producing H2 for a fuel cell (the

sweeping gas is in a co-current mode).

For performing a thermodynamic analysis and optimisation of the hydrogen-producing processes, a wide range of data is required. However, from literature, the experimental data related to the membrane reactors are normally at certain operation conditions. To fulfil the data requirement, the simulation of each unit, especially the simulation of the membrane reactors in the several possible processes has to be carried out firstly. With the simulation work as the basis, each process can then be designed, optimised and analysed, thus the most efficient process can be identified.

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_________________________________________________________________________________

1.4. Outline of the thesis

In Chapter 2, Data generation by simulation for process analysis and optimisation, the models for the simulation of the conventional CPO reactor,

membrane CPO reactors, conventional WGS reactor, and H2-membrane WGS reactor have been selected. Simulation methods have been described. With the simulation models, the axial temperature and the axial molar flow rates of different species are calculated. The simulation is the basis for further thermodynamic analysis. To test the validity of the model and the method, the simulation results have been compared with industrial and experimental data.

In the conventional CPO process and the membrane CPO processes, the CPO reactors play the most important role, because it is in the CPO reactors that methane is converted. Many factors, like the inlet temperature, pressure, inlet rate and ratio, and the configuration of the reactors have effects on how much inlet CH4 can be converted into the useful products (CO + H2), and how much CH4 should be combusted for providing the heat for the steam reforming reaction. Therefore, it is important to firstly have a look at the CPO reactors before comparing the integrated processes around these reactors. In Chapter 3, Comparison of

conventional and membrane CPO reactors, the effect of different inlet

conditions on the output of the CPO reactors have been investigated and then compared.

In Chapter 4, Thermodynamic analysis of integrated CPO processes with

constrained conditions, for the given inlet rates of reactants and the given

methane conversion, the thermodynamic analysis for each process with different adjustable inlet conditions is carried out. Corresponding to the different inlet conditions, different processes have been compared in terms of the thermodynamic efficiency, the production rate of H2, and the output of electric power. The most efficient process is indicated.

In Chapter 5, Thermodynamic analysis of an integrated process with

constrained geometry. The thermodynamic analysis of an integrated process with

a H2-membrane CPO reactor of fixed geometry for the production of separated H2 is described. The effect of inlet rate of methane, inlet ratio of methane to steam and the molar flow rate of sweeping gas on the production rate of separated H2 and the overall thermodynamic efficiency of the process is discussed.

As steam reforming process and CPO process are the two prominent processes for methane conversion, in Chapter 6, Brief analysis for conventional steam

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_________________________________________________________________________________ reforming (SR) process is described. Some results of the simulation and the thermodynamics analysis about the conventional SR process are presented.

The main results of this thesis as well as the main conclusions will be presented in Chapter 7, Summary of results and main conclusions. Chapter 8 is the Outlook.

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[4] Tsai C., Dixon A.G., Moser W.R., & Ma Y.H. (1997). Dense perovskite membrane reactors for partial oxidation of methane to syngas. AIChE Journal, Ceramic Processing, Vol.43, No. 11A, 2741-2750.

[5] Freni S., Calogero G and Cavallaro S. (2000). Hydrogen production from methane through catalytic partial oxidation reaction, Journal of Power Sources, 87, 28-38.

[6] Froment G.F. (2000). Production of synthesis gas by steam- and CO2- reforming

of natural gas, Journal of Molecular Catalysis, A, Chemical, 163, 147- 156. [7] Armor J.N. (1999). The multiple roles of catalysis in the production of H2,

Applied Catalysis: A 176, 159

[8] Momirlan M. and Veziroglu T.N. (2002). Currant status of Hydrogen energy, Renewable and Sustainable Energy reviews. 6, 141-179.

[9] Elnashair S. S. E. H. and Elshishini S. S. (1993). Simulation and Optimization of Industrial Fixed Bed Catalytic Reactors, Gordon and Breach, London, UK.

[10] Newsome D.S. (1980). The water-gas shift reaction. Catalysis Reviews- Science

and Engineering, 21(2), 275-318.

[11] Amphlett J.C., Mann R.F. and Peppley B.A. (1996) On board hydrogen

purification for steam reformation/ PEM fuel cell vehicle power plants, Int. J. Hydrogen Energy, 21(8), 673-678.

[12] Golunski S. (1998). HotspotTM_{Fuel Processor, Platinum Metals Review, 42(1),}

2-7.

[13] _{Avcl A.K., Trimm D. L. and O nsan Z. I. (2002). Quantitative investigation of}

catalytic natural gas conversion for hydrogen fuel cell applications. Chemical

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_________________________________________________________________________________

Engineering Journal, 4027, 1-11.

[14] Membrane Reactor Technology for CPO, Air Separation and Fuel Processing,

NWO project proposal, (1999).

[15] Ralph T. R. and Hograth M.P. (2002). Catalysis for low temperature fuelcells. Part

II. The anode challenges, Platinum Met. Rev. 46,117-135.

[16] Buxbaum R. and Lei H. (2003). Power output and load following in a fuel cell

fuelled by membrane reactor hydrogen, Journal of power sources, 123, 43-47.

[17] Hsieh H.P. (1991). Inorganic membrane reactors. Catalysis Reviews- Science and

Engineering, 33, 1-70.

[18] Paul N.D., Robin E.R., Steven L.R. and Dale M. T. (2000). Ion transport

membrane technology for oxygen separation and syngas production. Solid State Ionics, 134, 21-33.

[19] Kharton V. V., Naumovich E.N., Kovalevsky A.V., Viskup A.P., Figueiredo

F.M., Bashmakov I.A. and Marques F.M.B. (2000). Mixed electronic and ionic conductivity of LaCo(M)O3 (M = Ca, Cr, Fe or Ni) IV. Effect of preparation

method on oxygen transport in

LaCoO

₃₋_δ. Solid State Ionics, 138, 135-148.

[20] Maiya P.S., Balachandran U., Dusek J.T., Mieville R.L., Kleefisch M.S. and

Udovich C.A. (1997). Oxygen transport by oxygen potential gradient in dense ceramic oxide membranes. Solid State Ionic, 99, 1-7.

[21] Baker R.W. (2002). Future directions of membrane gas separation technology.

Industrial and Engineering Chemistry Research, 41, 1393-1411.

[22] Jin W., Li S., Huang P., Xu N., Shi J. and Lin Y.S. (2000). Tubular lanthanum

cobaltite perovskite-type membrane reactors for partial oxidation of methane to syngas. Journal of Membrane Science, 166, 13-22.

[23] Shao Z., Dong H., Xiong G., Cong Y. and Yang W. (2001). Performance of a

mixed-conducting ceramic membrane reactor with high oxygen permeability for methane conversion. Journal of Membrane Science, 183, 181-192.

[24] _{Balachandran U., Dusek J.T., Maiya P.S., Ma B., Mieville R.L., Kleefisch M.S.}

and Udovich C.A. (1997). Ceramic membrane reactor for converting methane to syngas. Catalysis Today, 36, 265-272.

[25] Jin W., Gu X., Li S., Huang P., Xu N., and Shi J. (2000). Experimental and

simulation study on a catalyst packed tubular dense membrane reactor for partial oxidation of methane to syngas. Chemical Engineering Science, 55, 2617-2625.

[26] Tsai C., Ma Y.H., Moser W.R. and Dixon A.G. (1995). Modeling and simulation

of a nonisothermal catalytic membrane reactor. Chemical Engineering Communications, Vol.134, 107-132.

[27] _{Galuszka J., Pandy R. N. and Ahmed S. (1998). Methane conversion to syngas in}

a palladium membrane reactor. Catysis Today, 46, 83-89

[28] Basile A., Paturzo L. and Lagana F. (2001). The partial oxidation of methane to

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Catalysis Today, 67, 65-75.

[29] Mohan K. and Govind R. (1988). Effect of Temperature on Equilibrium Shift in

Reactors with a Permselective Wall. Industrial and Engineering Chemistry Research, 27, 2064-2070.

[30] Amadeo N.E.and Laborde M.A. (1995). Hydrogen production from the

low-temperature water-gas-shift reaction: kinetics and simulation of the industrial reactor. International Journal of Hydrogen Energy,Vol.20(12), 949-956.

[31] _{Shah M. M., Drevich R. F., Balachandran U., Doris S.E. and Lee T.H. (2001).}

Technoeconomic feasibility analysis of hydrogen production by integrated ceramic membrane system, Proceedings of the 2001 DOE hydrogen program review. NREL/CP-570-30535.

[32] Chen Z., Yan Y. and Elnashaie Said S.E. H. (2003). Novel circulating fast

fluidised-bed membrane reformer for efficient production of hydrogen from steam reforming of methane. Chemical Engineering Science, 58, 4335-4349.

[33] Szargut J., Morris D.R. and Steward F.R. (1988). Exergy analysis of Thermal,

Chemical and Metallurgical processes. Hemisphere Publishing Corporation, New York.

[34] _{Ratkje S.K. and de Swaan Arons J. (1995). Denbigh revisited: reducing lost work}

in chemical processes. Chemical Engineering Science, 50(10), 1551-1560.

[35] Bejan A. (1996). Entropy generation minimization: The new thermodynamics of

finite-size devices and finite-time processes. Applied Physics Reviews, 79(3), 1191-1218.

[36] Springer T.E., Zawodzinski T.A. and Gottesfeld S. (1991) Polymer electrolyte fuel

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2. Data generation by simulation for process

analysis and optimisation

2.1. Introduction

For performing a thermodynamic analysis and optimisation of hydrogen-production processes, a wide range of data is required, because some important factors have to be taken into account, for example, pressure, temperature, membrane area, inlet molar ratio of methane to steam, etc. However, from open literature, the experimental data related to the membrane reactors are normally only available at certain operation conditions. To fulfil the wide range of data requirement, the possibility of simulation has to be explored, especially for the CPO and WGS reactors, without and with membranes.

Simulation of Conventional CPO and Conventional WGS reactors

Simulation of membrane reactors

H2-membrane CPO reactor H2-membrane WGS reactor O2-membrane CPO reactor Two-membrane CPO reactor

Fig.2.1. Simulation from conventional to membrane reactors

As shown in Fig. 2.1, because of the availability of experimental data for the conventional reactors, we will start with the simulation of the conventional reactors, and then compare the simulation results with the experimental data to prove the validity of the simulation work. By combining membrane permeation mechanisms, membrane CPO reactors will be simulated.

In the membrane reactors, the permeation rate of a component depends on the local partial pressure difference and temperature. The local partial pressure, temperature, the heat transferred by permeating species, the heat exchanged between reaction side and non-reaction side will surely affect the exit temperature and the compositions of each component in the product stream, and affect the chemical energy. So proper reaction kinetic models, membrane permeation

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__________________________________________________________________________________________ mechanism models and heat transfer models are the essential basis for obtaining more precise and reliable results.

In this chapter, the appropriate simulation models for the conventional catalytic partial oxidation (CPO) reactor and water-gas shift (WGS) reactors will be selected. The simulation results of the conventional reactors and membrane CPO reactors are compared with the experimental data to prove the validity of the simulation work. The simulation models and the simulation results for the membrane reactors will be discussed.

2.2. Simulation of the conventional CPO reactor

2.2.1. Kinetics

Among the different catalysts for the CPO reactor, transition metal catalysts, in particular Ni and Rh, are presently considered to be the most promising. The high-cost and limited availability of noble metals imply that Ni catalysts are preferred from an industrial standpoint [1]_.

For steam reforming of methane over a nickel supported catalyst, Xu and Froment [2]_{have suggested two main reactions for methane steam reforming}

together with a water gas shift reaction, and they have developed the corresponding kinetic rate equations, which are listed in Table 2.1. Elnashaie and Elshishini [3] reviewed a number of kinetic models available in the literature for steam reforming of methane over nickel-based catalysts and found that the kinetic models developed by Xu and Froment are the most general and reliable expressions for this process.

De Groote [4]_{simulated adiabatic fixed-bed reactors provided with a Ni catalyst}

for the catalytic partial oxidation of methane to synthesis gas at high temperatures and pressures. The intrinsic kinetics of the reforming, water-gas shift, and the oxidation reactions were taken from Xu and Froment[2]_{and Trimm and Lam}[5]_.

For the Boudouard reaction, methane cracking and carbon gasification by steam, the reaction kinetics were taken from unpublished results by Wager and Froment[6]_{. Special attention was paid to the calculations of coke formation and}

the influence of the degree of reduction of the catalyst. The influence of carbon dioxide and steam in the feed mixture on the coke formation rate was discussed. De Groote found that for feeds containing steam (steam to methane ratio above 1.5) there is almost no coke deposition.

With a simulation method similar to that of De Groote, Angelo Basile[7]_and

Ostrowski[8]_{simulated partial oxidation of methane to syngas in a H}

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__________________________________________________________________________________________ reactor. The influence of carbon deposition due to the Boudouard reaction and that of the cracking of methane on the catalyst activity were neglected. The simulation results were in a good agreement with the experimental data.

In this work, the simulation of the CPO reactors is based on the kinetics of the methane combustion reaction, the kinetics of the steam reforming reactions to CO and CO2, and on the kinetics of the water-gas-shift reaction chosen by de Groote [4]_{. The rate equations as well as the kinetic parameters applied in the calculations}

of the reaction rate are summarized in Tables 2.1 –2.3.

In these tables,

R

_k, , and represent the reaction rate, the reaction rate constant, the pre-exponential factor for the reaction rate constant and the equilibrium constant of reaction , respectively.

k

0

k

K

_eqk

k

p

_i, and are the partial

pressure, the adsorption constant, and the pre-exponential factor for the adsorption constant of species

i

, respectively. is the activation energy of reaction .

is the heat of adsorption of species on the catalyst, respectively.

i

K

0 i

K

k i ads

H

_,

∆

k a,

E

i

With the aid of a thermodynamic analysis using AspenPlus, Seo[9]_{examined the}

feed ratios of air and steam to methane. The results showed that no carbon formation was observed as the feeding ratios of air and steam to methane are over 0.3 and 1.0, respectively. Based on the results of De Groote[4], Basile[7], Ostrowski[8]_{and Seo}[9]_{, the influence of carbon deposition due to the Boundouard}

reaction and that of the cracking of methane on the catalyst is neglected (in this work, for the CPO reactors, the inlet ratio of steam to methane is above 1.0, the inlet ratio of air to methane is above 0.3).

Table 2.1. Reaction rate for the reactions in the CPO reactors.

Reactions Kinetics of reaction

1 CH4+2O2UCO2+2H O2 4 2 4 2 4 4 2 2 4 4 2 2 1/ 2 1, 1, 1 ₍₁ 0 0 ₎2 ₍₁ 0 0 ₎ a CH O b CH O CH CH O O CH CH O O k x x k x x R K x K x K x K x = + + + + + 2 CH4+H O2 UCO+3H2 2 4 2 2 2.5 3 2 2 2 2 ( /k pH)(pCH pH O p pH CO/Keq ) R DEN − = 3 CO+H O2 UH2+CO2 2 2 2 2 3 3 3 2 ( /k p_H)(p p_{CO H O} p p_H _CO/K_eq) R DEN − = 4 CH4+2H O2 UCO2+4H2 2 4 2 2 2 3.5 2 4 4 4 4 2 ( /k p_H)(p p_CH _{H O} p p_H _CO/K_eq) R DEN − = Note: DEN = +1 K pCO CO+K pH₂ H₂ +KCH₄pCH₄ +KH O₂ pH O₂ /pH₂

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__________________________________________________________________________________________ Table 2.2. Parameter values for the expression _k0exp( _E _, _RT)

k a k

k = −

k for the

reaction rate constant.

Reaction

_k

0

k[mol kg/( cal⋅s)]

E

a,k [J mol/ ] 1

The transport mechanism in the axial direction is considered to be of the plug flow type. The influence of intra-particle concentration gradients within the catalyst pellet is taken into account by solving the solid phase continuity equation at each increment along the adiabatic fixed bed reactor coordinate.

The gas-phase continuity equation (or mass balance equation), energy equation (or energy balance equation), and solid phase continuity equation are presented in Table 2.4, corresponding inlet and boundary conditions are also listed. In the

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__________________________________________________________________________________________ reactor model, the subscript

i

represents the reaction gas species, i.e., CH4, H2O,

CO, H2, CO2, O2 and N2. k

In Table 2.4,

η

and are the effectiveness factor and heat of reaction , respectively.

k k

ik

H

∆

υ

represents the stoichiometric coefficient of component taking part in reaction .

i

k

ε

_B,

ρ

_s, , and z represent the void fraction of the catalyst packing, catalyst density, cross sectional area and axial coordinate of the reactor, respectively. and

r A

i

F

p

_{s i}_, are the molar flow rate and partial pressure of component inside the catalyst particle, respectively.

i

ξ

represents the dimensionless particle

radius. The area of the catalyst particle per unit mass and the equivalent radius are represented by

a

_vand

r

_s, respectively.

In the solid phase continuity equation, the effective diffusivity of species ( ) is related to the molecular ( ) and Knudsen diffusivities ( ) by Eq. (2.1), in which

S r

)

(

)

(2.2)

where

S r

is the void volume fraction taken of a pore with radius obtained from Xu [10]_{. Using the average value of diffusivity} _{, an effective}

diffusivity is calculated through Eq. (2.3)

pore

r

pore i D , e i s i

D

ε

D

τ

=

_(2.3)

The tortuosity factor τ is the simplified and lumped effect of many factors. is the porosity of the catalyst particle. In the simulation, (0.53) and τ (4.0) have

s

ε

s

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__________________________________________________________________________________________

i

p

C

been obtained from literature [10]_{. The physical chemical properties} [11]_,

equilibrium constants [12]_{, and diffusivities}[13]_{are considered to be temperature}

dependent.

2.2.3. Solution procedure

The model for a conventional CPO reactor consists of a set of differential equations. The solid phase equations are solved by a difference method. The number of collocation points, which is to be defined, depends on its effect on the final calculation results. Normally 10 collocation points are defined. The resulting difference equations are solved at each increment of the axial reactor coordinate. In each reactor increment, the intra-particle concentration gradients of the previous step are used as the initial value of the next step to solve the solid-phase continuity equations. In this way, rapid convergence is obtained.

Table 2.4. Conventional CPO reactor model Gas phase continuity equation

1 (1 ) R s N c p o i k ik k B r k d F R a d z = η υ ε ρ =

∑

−

Gas phase energy equation 7 (1 ) ( ) R s i cpo N B r k k k k i p i a dT H R dz _{F C} ρ ε η − =

∑

− ∆

∑

Solid phase equations for calculating the effectiveness factors , 2 5 2 , , 2 1 , , 5 1 1 ( ) 10 ( 10 ( ) R R N s i cpo e i s s i k s k k N e i s i v k ik k k s dp d D RT r d d D dp a RT R r d ξ ξ ρ ν ξ ξ ξ η υ ξ − = − = = =

∑

, ) R

Solid phase boundary conditions , 0,dps i d ξ ξ 0 = =

Gas phase boundary conditions

0

0,

_i _{i z}

,

_in

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__________________________________________________________________________________________

2.2.4. Oxidation state of the catalyst

For the simulation of an industrial autothermal reformer for syngas production, De Groote [4]_{used two models for simulating the oxidation states of the Ni}

containing catalysts. In the first model, the catalyst has a varying degree of reduction (the VDR-model) and the steam reforming is mainly consecutive to the combustion of methane. In the second model the catalyst is bivalent (BV-model) and the catalytic combustion and steam reforming operate purely in parallel. De Groote incorporated the effect of the variable Ni oxidation state by multiplying the rates of the steam reforming and water gas shift reactions with a so-called reduction factor. The reduction factor is assumed to depend on the fractional oxygen conversion according to ( . The power in this reduction factor was determined with respect to the calculated maximum catalyst temperature, using a CH 2 12 ) O X 4/O2 feed.

With the same reduction factor as De Groote[4]_{has used, the temperature axial}

profiles calculated with the VDR-model and the BV-model in this work are shown in Fig. 2.2. In the case of the VDR-model, the catalyst temperature first increases due to the large amount of heat produced in the methane oxidation reaction. The temperature reaches a maximum value, and then decreases gradually as a result of the prevailing endothermic reforming reaction. Due to the complex interplay between the heat consumed by the reforming reaction, and the heat produced in the water-gas shift and oxidation reactions, the simulation results indicate a shallow minimum temperature. The temperature finally increases to the equilibrium temperature. With the BV-model, the temperature initially decreases due to the prevailing endothermic reforming reactions. After the temperature reaches a minimum value, it subsequently increases to the equilibrium temperature.

Figs. 2.3a and 2.3b illustrate the axial profiles of the molar fractions for different species calculated with the BV model and the VDR model, respectively. As shown in Figs. 2.3a and 2.3b, both methane and oxygen molar fractions decrease rapidly at the beginning of the reactor. Significant amounts of H2 and CO are

observed already in the first part of the reactor. In the second part of the reactor, the molar fractions calculated by both models are almost constant as the equilibrium conditions are being approached.

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__________________________________________________________________________________________ 0.0 0.2 0.4 0.6 0.8 1.0 700 800 900 1000 1100 1200 1300 1400 VDR-model BV-model T (K ) Axial coordinate [m]

Fig. 2.2. The temperature vs. axial coordinate of the conventional CPO reactor (without membrane) at Tin= 800 K and

P

_in =10 bar. The inlet rates of CH4, H2O and O2 are 1.0, 1.5

and 0.535mol/s, respectively.

0.0 0.2 0.4 0.6 0.8 1.0 0.00 0.05 0.10 0.15 0.20 0.25 0.30 0.35 0.40 BV-model a O₂ CH₄ CO₂ CO H₂ H₂O M o la r f ac tio n [ -]

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__________________________________________________________________________________________ 0.0 0.2 0.4 0.6 0.8 1.0 0.00 0.05 0.10 0.15 0.20 0.25 0.30 0.35 0.40 VDR-model b O₂ CH₄ CO2 CO H₂O H₂ M o la r fra cti on [ -]

Dimensionless reactor axial coordinate [-]

Fig. 2.3. The molar fractions of different species simulated by the BV-model (a) and by the VDR-model (b) vs. the dimensionless axial coordinate of the conventional CPO reactor at Tin= 800 K.

P

in =10 bar; the inlet rates of CH4, H2O and O2 are 1.0, 1.5 and

0.535mol/s, respectively.

2.2.5. Validity of the conventional CPO reactor model

The reforming model proposed by Xu and Froment was obtained using relatively low temperature (773 K to 848 K) and pressures between 3 and 15 bar. In a CPO reactor for methane conversion, a higher temperature can occur and a higher pressure may be required. To prove the validity of the simulation model, the simulation results (the VDR model) of the conventional CPO reactor are compared with industrial data at an inlet temperature of 808 K and an inlet pressure of 25 bar[4]_{. The results are shown in Table 2.5a. It is clear that the result} at the exit is in a good agreement with those measured in industrial practice, when the pressure and the temperature are beyond the range in which Xu and Froment proposed the reforming model.

It should be noted that the kinetic rate constants for methane oxidation obtained by Trimm and Lam is for a Pt catalyst. De Smet [1]_{adjusted the kinetic rate} constants for methane oxidation on Ni catalyst through many assumptions, but De Smet et al have not checked their simulation results with experimental data or industrial data. Table 2.5b shows the simulation results of De Smet and the simulation results by our simulation program with the methane oxidation model by Trimm and Lam. With the same geometric values and the same inlet condition

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__________________________________________________________________________________________ for the CPO reactor, our simulation results at the exit of the reactor are very close to the results calculated by De Smet. The difference in maximum temperatures is within 30 K. Because the limited influence of the kinetic rate constants for the methane oxidation on the simulation results, in this work the same kinetic rate constants obtained by Trimm and Lam are used, as listed in Table 2.2.

The comparison results shown in Tables 2.5a and 2.5b prove that the model can be used at a higher pressure.

Table 2.5a. Comparison of the simulation results with the industrial data [4]_{for an}

autothermal reformer.

Industrial Simulation

Feed composition and conditions F0 CH4(mol/s) 40.24 40.24 O2/CH4 0.598 0.598 H2O/CH4 1.4 1.4 in

P

in

T

(bar) 25.33 25.33 (K) 808 808

Product yield (molar fraction)

yO2 0 0.001 yCH4 0.008 0.007 yH2 0.456 0.476 yCO 0.160 0.162 yCO2 0.070 0.062 yH2O 0.306 0.292 Outlet conditions Tout(K) 1223 1243 FH2(mol/s) 80.61 82.97 FCO(mol/s) 28.24 28.17

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__________________________________________________________________________________________

Table 2.5b. Comparison of the simulation results of this work with the simulated results by De Smet[1]_{for a conventional CPO reactor.}

De Smet This work

Feed composition and conditions F0 CH4(mol/s) 96.1 96.1 CH4/ O2 1.8 1.8 H2O/CH4 1.0 1.0 in

P

in

T

(bar) 40.0 40.0 (K) 773 773

Product yield (molar fraction)

H2/CO 2.64 2.61

CH4 conversion 94.6 94.6

Tmax (K) 1510 1540

Tout(K) 1287 1287

Note: The reactor geometric parameters are the same as that listed in the literature [1]_:

reactor diameter, 1.6m; reactor length, 3m; void of packing, 0.43; pellet radius, m.

3

7.5 10_× −

2.3. Simulation of membrane CPO reactors

2.3.1. Reactor model for membrane CPO reactors

The simulation model for the membrane CPO reactors is shown in Table 2.6. We have assumed plug flow at both sides of the membrane. The gas-phase continuity, energy, solid continuity equations on the reaction side as well as the continuity and energy equations on the non-reaction side are presented in Table 2.6, in which the corresponding initial and boundary conditions are also listed.

The energy equations in Table 2.6 take into account the heat of reaction, the heat exchanged between the non-reaction zone and the reaction zone, and the energy carried by the diffusing O2 or H2.

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__________________________________________________________________________________________

Table 2.6. The reactor model for membrane CPO reactors

Reaction side

Gas phase Continuity equation

2 2 2 2 2 2 2 2 2 2 1 4 6 (1 ) ( ) ; 0, 4; ; 0, 6 R N H O cpo i r B s k ik k H i O i k H H O O H i O i dF a R a N a dz N N N i N N N i ε ρ η υ = = − − + = = ≠ = = ≠

∑

N

Gas phase energy equation

2 2 2 2 2 2 2 2 7 1 1 1 [ (1 ) ( ) ] ( ), ( ) R i N CPO CPO r r s B k k k H H H O O O k i p i

H m CPO CPO O m CPO CPO

r nr r nr sp sp dT a H R q a N H a N H dz F C a k a k q T T or q T T ρ ε η δ δ = = = − −∆ − − ∆ + ∆ = − = −

∑

Solid phase equations for calculating the effectiveness factors

, 2 5 2 , , 2 1 , , 5 1 1 ( ) 10 ( 10 ( ) R R N s i CPO e i s s i k s k k N e i s i v k ik k k s dp d D RT r d d D dp a RT R r d ξ ξ ρ ν ξ ξ ξ η υ ξ − = − = = =

∑

, ) R

Non-reaction side Continuity equation 2 2 2 2 2 H O 2 H H O dG dG a N or a N dz = dz =− O Energy equation 2 2 2 2 2 2 1 [ ] j CPO nr H H H O O O j p j dT q a N H a N H dz G C = + ∆ − ∆

∑

Solid phase boundary conditions , 0,dps i d ξ ξ 0 = =

Gas phase boundary conditions 0

0, i i z ,

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__________________________________________________________________________________________ 2

In Table 2.6, and represent the permeation rate of component

i

through a H

2-membrane area and

the O2-membrane area per unit reactor length, respectively. and represent the heat transferred by the permeation of H

2

O

H

∆

2 or O2 between the

reaction side and the non-reaction side, respectively.

q

represents the heat exchanged between the reaction side and non-reaction side. The molar flow rate of different species

j

in the non-reaction side is represented by

G

.

H i

N

2 2 H

H

∆

j

Both the O2-membrane and the H2-membrane consist of two layers (see Fig.2.4), a

support layer with large pores and a dense membrane layer. The thickness of the

support layer ( ) is assumed to be . Because of the lower

permeability of the membrane

0.005 m

[14,15]_{, the thickness of the dense membrane (} _or

) is assumed to be

5 µ

m

in Chapters 3 and 4 to increase the O2 and H2

permeation rate. (For a given inlet rate of methane and a given methane conversion, this assumption does not affect the production rate of useful products (CO + H2) of membrane reactors.)

k

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__________________________________________________________________________________________ 1 1 2 m sp k =k +

δ

_{sp s} y k

y

(2.4)

where represents the thickness from the centre of the catalyst bed to the membrane wall. and

δ

are the thermal conductivity and thickness of support layer. For the membrane CPO reactor, the thermal conductivity

k

_sof the porous catalyst Ni-Al2O3 is about 0.3

J m s K

/(

⋅ ⋅

)

[16]. Since the thermal conductivity of the support layer, such as ceramic Al2O3 is about 8

J m s K

/(

⋅ ⋅

)

at around 1000K [17]_{, which is much larger than that of the catalyst bed, the heat transfer resistance} mainly comes from the packed bed. Due to the low permeability of the membrane, a higher ratio of membrane area to reactor volume is required,

/

y

δ

is around 2 - 4 in this work. sp

k

_sp

sp

In the simulation, the composition of the perovskite membrane for O2

permeation is

La Ba Fe Co O

_0.2 _0.8 _0.8 _0.2 ₃−δ. The permeability of oxygen through this

membrane is calculated from Eq.(2.5).

2 2 2 2 / 1 a

_ln(

₎

high E RT o O low O o

p

A e

T

N

p

δ

−

=

(2.5)

The activation energy and the pre-exponential factor are

63

and , respectively. The oxygen permeation rate equation, and the coefficients have been obtained from literature

Ea A₁

kJ

) ⋅ ⋅

/ mol

7 7.34 10 mo_× − _{l m}_/( s K [14,18]_.

The permeability of hydrogen through the H2 permeable membrane is calculated

according to Eq.(2.6) 2 2 2

exp(

)

(

)

A m high low H H H

E

P

RT

N

p

δ

−

=

−

2 H

p

(2.6)