gasification of solid fuels: countercurrent reactor...

Gasification of Solid Fuels:

Countercurrent Reactor Modeling

and Metallic Iron Use for Tar

Removal

MARIAROSA BRUNDU

DOTTORATO DI RICERCA IN INGEGNERIA INDUSTRIALE UNIVERSITÀ DEGLI STUDI DI CAGLIARI XXII CICLO

Gasification of Solid Fuels:

Countercurrent Reactor Modeling

and Metallic Iron Use for Tar

Removal

MARIAROSA BRUNDU

Supervisor:

Professor Giampaolo Mura

Dottorato di Ricerca in Ingegneria Industriale

Università degli Studi di Cagliari

XXII Ciclo

“Denique consilium rerum omnium sapiens,

non exitum spectat; initia in potestate nostra sunt,

de eventu fortuna iudicat, cui de me sententiam

non do.”

Lucio Anneo Seneca

…

…

…

“To the future or to the past, to a time when

thought is free, when men are different from one

another and do not live alone, to a time when

truth exists and what is done cannot be undone.

From the age of uniformity, from the age of

solitude, from the age of big brother, from the age

of doublethink

Greetings!”

George Orwell, 1984

…

…

…

“I know I was born and I know that I’ll die,

the in between is mine”

Pearl Jam, I am mine

i

Contents

Abstract .......................................................................................................1

Introduction ............................................................................................... 3 A brief overview on the role of coal and biomass in the world energy production ............................................................................................................. 5

1.1 The World Energy consumption: a prediction up to 2030 .................. 7 1.2 Recent developments on the European energy policy ......................... 8 1.3 Biomass and coal as a competitive solution for a clean energy production ......................................................................................................... 8

1.3.1 Biomass potential ............................................................................... 9 1.3.2 Coal potential .................................................................................... 10 1.3.3 A brief summary of coal and biomass applications .................... 11

1.4 References ................................................................................................. 12 Fundamentals of the gasification process of solid fuels ............................... 13

2.1 Introduction .............................................................................................. 15 2.2 An overview on the gasification process .............................................. 17

2.2.1 Feed pretreatment ............................................................................ 17 2.2.2 Reaction step and reactor types ..................................................... 17 2.2.3 Gas clean up section ........................................................................ 21 2.2.4 Influence of parameters................................................................... 24

2.3 References ................................................................................................. 26

A mathematical model for a fixed bed updraft reactor ........................... 29 Nomenclature ................................................................................................. 31 Introduction .................................................................................................... 33 References ....................................................................................................... 35

The model ................................................................................................................ 37 Model description .............................................................................................. 39

3.1 Introduction .............................................................................................. 41 3.2 Governing equations ............................................................................... 41

3.2.1 Mass balance ..................................................................................... 41 3.2.2 Heat balance ...................................................................................... 43 3.2.3 Other assumptions ........................................................................... 44

3.3 Kinetic description ................................................................................... 47 3.3.1 Drying ................................................................................................ 47 3.3.2 Pyrolisys ............................................................................................. 48

3.3.2.1 Coal ............................................................................................. 48 3.3.2.2 Biomass ...................................................................................... 50

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3.3.3 Combustion and gasification .......................................................... 50 3.4 References ................................................................................................. 55

Model results ........................................................................................................... 61 Introduction .................................................................................................... 63 Numerical solution......................................................................................... 63 Initial conditions............................................................................................. 66 References ....................................................................................................... 67

Biomass gasification .......................................................................................... 69 4.1 Introduction.............................................................................................. 71 4.2 Dynamic behavior ................................................................................... 72 4.3 Steady state description ........................................................................... 80 4.4 Sensitivity analysis .................................................................................... 85

4.4.1 Influence of the Air To Fuel ratio (ATF) ..................................... 85 4.4.2 Influence of the Steam To Fuel ratio (STF) ................................. 87 4.4.3 Influence of gasification reactions ................................................. 89 4.4.4 Influence of the gas-solid heat exchange ...................................... 90 4.4.5 Influence of the initial bed height .................................................. 91

4.5 Conclusions .............................................................................................. 92 4.6 References ................................................................................................. 93

Coal gasification ................................................................................................. 95 5.1 Introduction.............................................................................................. 97 5.2 Dynamic behavior ................................................................................... 98 5.3 Steady state description ......................................................................... 101 5.4 Sensitivity analysis .................................................................................. 105

5.4.1 Influence of the Air To Fuel ratio (ATF) ................................... 106 5.4.2 Influence of the Steam To Fuel ratio (STF) ............................... 108 5.4.3 Influence of the solid flowrate ..................................................... 110 5.4.4 Influence of the reactor length ..................................................... 110

5.5 Conclusions ............................................................................................ 113 5.6 References ............................................................................................... 114

Gas clean up: tar removal ........................................................................ 117 Iron material exploitation for biomass tar removal: an experimental study ............................................................................................................................ 119

6.1 Introduction............................................................................................ 121 6.2 Tar sampling and analysis ..................................................................... 124 6.3 Tar removal ............................................................................................ 126

6.3.1 Physical removal ............................................................................. 126 6.3.2 Chemical removal........................................................................... 127

6.3.3.1 Thermal reduction .................................................................. 128 6.3.3.2 Steam Reforming .................................................................... 128 6.3.3.3 Catalytic cracking .................................................................... 128

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6.4 Experimental setup ................................................................................ 130 6.4.1 Material characterization ............................................................... 130 6.4.2. Apparatus ....................................................................................... 132 6.4.3 Experimental performance ........................................................... 133 6.4.4 Gas and tar analysis ........................................................................ 135

6.5 Results and discussion ........................................................................... 136 6.5.1 Catalyst performance and tar conversion ................................... 137 6.5.2. Influence on gas composition ..................................................... 138 6.5.3 Catalyst Life ..................................................................................... 138

6.6 Conclusions............................................................................................. 139 6.7 References ............................................................................................... 140

Conclusions ............................................................................................. 143

1

Abstract

This work deals with the gasification of solid fuels. It is divided in two sections. The first deals with the development of a mathematical model of the transient behavior of a countercurrent fixed bed reactor for coal and biomass gasification. This section is organized in two parts: a detailed description of the model and model results for both, a biomass feedstock and a coal feedstock. The one dimensional model developed in this study is based on heat and mass continuity equations. All the main phenomena involved in the gasification process are inserted: drying, pyrolysis, gasification and combustion reactions of the solid phase and homogeneous gas phase reactions. A kinetic approach is chosen in this work to address the description of gasification reactions: global kinetic models are used for pyrolysis; heterogeneousness of the system is considered for gas phase reactions. Two separate heat balances are inserted: one for the gas phase and one for the solid phase. Differently from previous works, the ash content of the fuel is considered and a mass balance on this component is inserted in the model.

Results concerning the transient behavior of an air blown atmospheric gasifier are presented in the second part of the section. Results are organized in two chapters respectively referring to a birch wood feedstock and to a Pittsburg n°8 coal seam feedstock. Every chapter starts with the description of the transient behavior of the reactor. In order to define the initial conditions the start up procedure of an existing coal gasification plant is chosen as a reference. Since a look to the literature showed scarce information concerning long term dynamics of the reactor a description of the bed behavior and its change in composition with time is given. Composition and temperature profiles along the reactor at steady state condition for both, the gas and the solid phase are discussed. A sensitivity analysis of the system is presented at the end of every chapter. The influence of the main parameters affecting the system is examined.

The second section deals with gas clean up. One of the major problems encountered during biomass gasification is, indeed, the tar content of the gas product. Physical or chemical methods are usually used to destroy these undesired components. The exploitation of three types of iron material as catalyst for tar cracking was studied. Materials, provided by Höganäs AB, were tested with real raw gas coming from the atmospheric air blown fluidized bed

2

gasifier existing at KTH (Stockholm). The system used for investigation consists of a gasification reactor, a filter for particle removal and a catalytic reactor for tar conversion. Each sample was tested at 700, 750, 800, 850 °C. Birch wood is the feedstock used to run the simulations. Results show the influence of temperature on catalyst activity and comparisons between the three materials.

Introduction

Chapter

1

A brief overview on the role of coal and

biomass in the world energy production The role of coal and biomass in the increasing world energy consumption over the years is presented in this Chapter. In particular the need to develop new technologies and to improve existing ones in order to satisfy the new European energy policy is pointed out

CHAPTER 1. THE ROLE OF COAL AND BIOMASS IN THE WORLD ENERGY PRODUCTION

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1.1 The World Energy consumption: a prediction up to 2030

According to the IEO2009 projections the total energy consumption is projected to increase by 44% from 2006 to 2030; the largest increase is expected for the non-OECD economies [1].

Given expectations that world oil prices will remain relatively high through most of the projection period, liquid fuels are the world’s slowest growing energy source while renewables are the fastest with consumption increasing by 3% per year. From 2006 to 2030 liquids consumption in the residential, commercial and electric power sectors are, therefore, expected to decline on a worldwide basis.

Since it is more efficient, less carbon intensive than other fossil fuels and cheaper than oil natural gas remains an important energy source for electricity generation worldwide. According to the prevision the world coal consumption is expected to increase by 1.7% per year on average from 2006 to 2030 and to account for 28% of total world energy consumption in 2030. In the absence of policies or legislation that would limit the growth of coal use, the United States, China, and India are expected to turn to coal in place of more expensive fuels. The only decreases in coal consumption are projected for OECD Europe and for Japan where renewable energy sources, natural gas, and nuclear power are likely to be chosen over coal for electricity generation. Much of the growth for renewable energy sources is in hydroelectric power and wind power. As renewable energy use increases worldwide, the mix of fuels in the OECD and non-OECD regions differs. In the OECD nations, the majority of economically exploitable hydroelectric resources have already been developed. Instead, most renewable energy growth in the OECD countries is expected to come from non hydroelectric sources, especially wind and biomass. In contrast to the OECD countries, hydroelectric power is expected to be the predominant source of renewable energy growth in the non-OECD nations.

The major uncertainties within this projection are expectations for the future rates of economic growth and oil price. However the projections for total world energy consumption in 2030 do not vary substantially while the distribution between coal, natural gas and oil strongly depends on the oil price. The difference for renewable energy consumption is instead very small because their development mainly depends on government policies and incentives more than

INTRODUCTION

8

fuel price. In order to face the energy production scenario the need to develop and improve biomass based technologies is worldwide recognized.

1.2 Recent developments on the European energy policy

In the past the Kyoto protocol was the reference guideline for the reduction of greenhouse gas emissions all over the world. A new reference document is nowadays available for the EU. During the meeting held in Brussels on January 23rd,2008 a new strategy for energy production with relation to the limitation of pollutants emission was presented by the EC council. Two key targets were set by at that time [4]:

� A reduction of at least 20% in greenhouse gases (GHG) by 2020 – rising to 30% if there is an international agreement committing other developed countries to "comparable emission reductions and economically more advanced developing countries to contributing adequately according to their responsibilities and respective capabilities".

� A 20% share of renewable energies in EU energy consumption by 2020.

The goal of the European Union is very ambitious and a close interaction between different fields is necessary to achieve it: politics, science and economics. A great amount of money and resources is necessary especially to improve the existing technologies for fossil fuels and to develop new technologies for the exploitation of renewable energy sources. The improvement of energy production efficiency also plays a key role on the reduction of emissions.

1.3 Biomass and coal as a competitive solution for a clean

energy production

A definition of biomass and coal together with some properties and classifications is given in the following, moreover, the available and under development technologies for their exploitation is discussed, particular attention is paid to their potential as energy sources.


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1.3.1 Biomass potential

Biomass is every biological material derived from living or recently living organisms: plants, animals and waste. Forest residues (e.g. dead trees, branches and tree stumps), yard clippings and wood chips may be referred as biomass. However, biomass also includes plant or animal matter and biodegradable wastes. Biomass is carbon based and is composed by a mixture of organic molecules containing hydrogen, oxygen, often nitrogen and also small quantities of other atoms including alkali, alkaline earth and heavy metals. Biomass is usually characterized by proximate and ultimate analysis. The volatile content, fixed carbon content, moisture and ash content is given in the first; the percentage of some elements composing the fuel (usually C, O, H, N and S) is given in the second. A strong relation between biomass composition and the process where this kind of biomass is used exist. Moisture and nitrogen content are for example important in order to choose between thermal or biological treatments. Biomass as Ligno-cellulosic material is considered in this work.

With declining reserves, fluctuating prices of fossil fuels and the need of a cleaner policy for energy production, the search for an alternate renewable raw material to replace petroleum has been intensified all over the world. Because of its wide spread availability, renewable in nature and neutral in relation to global warming, one of the most important energy sources amongst the renewable energies in the near future is biomass.

Justification for bioenergy includes: security of long-term energy supplies for Europe, contributions to the development of industrial markets, improvements of the environment by utilizing wastes and residues, making a positive contribution to limit the greenhouse effect, better management of surplus agricultural and marginal land, provision of opportunities for socio-economic development of the less-developed regions of Europe, particularly towards the south.

One of the major problems with biomass is that, as an energy crop, it is labor-intensive to produce, harvest and transport, it is dispersed over large areas and can affect the price of food. When it is in the form of wastes, costs are much lower, often negative in the case of domestic solid waste, but material usually requires extensive processing to make it compatible with the conversion process. As produced, biomass is a solid and is difficult to use in many applications without substantial modification. Bio-chemical and thermo-chemical processes are used for energy recovery from biomass and conversion

INTRODUCTION

10

to gaseous and liquid energy carriers (from hydrogen to liquid transport fuels) that have many advantages in handling and application

1.3.2 Coal potential

Coal is a heterogeneous organic fossil fuel formed from vegetable matter by the action of heat and pressure during ages. It is composed by organic and inorganic matter. Coal composition strictly depends on the original organic matter and conditions during its formation process; a wide distribution of coal types exists. Like biomass they are usually characterized by a proximate and an ultimate analysis. Many classification of coal type are available in the literature: they are based on proximate and ultimate analysis, heating value and petrographic components. The most important coal ranking classifies the different types into four main series: lignite, sub bituminous, bituminous and anthracite. More sub series exist and further classification is possible. Depending on this coal ranking some considerations about its properties are possible: lignite has the lower carbon content and higher oxygen content; anthracite has the higher carbon content and the lower oxygen content. A high amount of volatiles are released from lignite if compared to anthracite, the latter has the highest calorific value. Nevertheless, this coal is less reactive. Surface area, an important parameter for all the thermal processes, is higher for lignite than for anthracite and in general decreases with the carbon content of the fuel.

The possibility to cover the overall energy demand with biomass is far from being realistic and the interaction of many sources is necessary. Among the different available energy sources coal can still have an important role but its clean use has to be promoted. In the past it was considered the dirtiest fuel and its utilization was connected to high emissions. A lot of work has been done in the last decades and with new technologies, especially with the option of carbon capture and storage, coal importance is nowadays increasing. A clean use of coal is a cheap alternative to oil and a co-utilization with biomass is often considered suitable by many researchers. Indeed coal can be used to make up for some biomass negative aspects; in fact, assuring a continuous run of power plants the problem of seasonality or wide storage of the fuel can be avoided and the need of small plants limited.


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1.3.3 A brief summary of coal and biomass applications

A wide range of processes for biomass conversion into more valuable fuels and/or energy is nowadays available. Among these we can distinguish: biological processes, usually aimed at ethanol or methane production; and thermal processes in order to produce heat, gaseous fuels, liquid fuels and solid fuels. A wide variety of secondary products, including electricity, can be produced starting from coal.

Electricity by thermal conversion of biomass is achievable through several routes: pressurized gasification plus a gas turbine in a combined cycle mode (IGCC technology is relatively well developed, with many systems commercially available), atmospheric gasification plus a turbine or an engine and simple combustion.

The most used application for direct combustion of biomass is heat production. The use of firewood for heat generation in order to keep houses warm is the most common application especially in the poorest countries. Nevertheless, direct combustion of solid biomass is sometimes used to generate electricity. A number of steam cycle power plants based on agricultural and forest industry residues have been built around the world, however, at a scale suited to biomass applications, such plants are expensive and have a low efficiency (about 32%); their use is usually connected to low feedstock costs and high electricity prices [2]. Complete combustion is a complex phenomenon, it proceeds in many intermediate steps such as drying, pyrolysis and combustion and it requires sufficiently high temperature, strong turbulence of the air-gas mixture, and a long residence time of the mixture in the reactor.

The main coal applications are: pulverized coal combustion, atmospheric or pressurized fluid bed of fixed bed combustion and integrated gasification combined cycle. Other applications for coal use are: carbonization, cooking and liquefaction through a pyrolysis or extraction step. Electric power generation is expected to be the principal use of both, coal and biomass, in the near- or mid- term; the production of liquid and gaseous fuels starting from them is on the contrary expected to compete with oil and natural gas in the mid and long term, especially in case of a large increase of the oil cost or its shortage. In particular, the application of the gasification process to biomass and coal seems to be very promising. Higher conversion efficiencies are achievable with this process and even the gas cleaning seems to be easier in comparison with simple combustion technologies.

INTRODUCTION

12

1.4 References

1. AA.VV., International energy outlook 2009, U.S. Department of Energy report DOE/EIA-0484(2009), Washington, 2009. (Available on line at www.eia.doe.gov/oiaf/ieo/index.html)

2. Bridgwater A. V., The technical and economic feasibility of biomass gasification for power generation, Fuel, 74, 1995.

3. Cipollina A., Lo Verso G., Micale G., L’utilizzo del carbone come combustibile economico e di limitato impatto ambientale, Terzo convegno internazionale Energia e Ambiente, Sorrento, Campania, Italy, 2004

4. Commission of the European Communities, 20 20 by 2020 Europe's climate change opportunity. COM(2008) 30 final, Brussels, 23.1.2008

5. Grimston M. C., Coal as an energy source, IEA Coal Research (CCC/20), 1999.

6. Küçük M. M., Demirbaş A., Biomass conversion process, Energy Conversion and Management, 38, 1997,151-165.

7. Longwell J. P., Rubin E. S., Wilson J., Coal: energy for the future, Progress in Energy Combustion Science, 21, 1995, 269-360.

8. Qader S. A., Coal science and technology - Natural gas substitutes from coal and oil, Amsterdam, 1985 Elsevier.

9. Saxena R. C., Seal D., Kumar S., Goyal H. B., Thermochemical routes for hydrogen rich gas from biomass: a review, Renewable and Sustainable Energy Reviews, 12, 2008, 1909-1927.

10. Smith I. M., Rousaki K., Prospects for co-utilization of coal with other fuels – GHG emissions reduction, IEA Coal Research (CCC/60), 2002.

Chapter

2

Fundamentals of the gasification

process of solid fuels In this Chapter the fundamentals of gasification of solid fuels are summarized. An overview of the process, its major steps, type of reactors and of the most important influencing parameters is given. A complete review of the gasification process is out of the scope of this work and is not addressed here but more complete information is

available in the cited literature.

CHAPTER 2 FUNDAMENTALS OF THE GASIFICATION PROCESS OF SOLID FUELS

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2.1 Introduction

Thermochemical gasification is the conversion by partial oxidation at elevated temperature of a carbonaceous feedstock (e.g. biomass or coal) into a gaseous energy carrier.

The use of this process for coal conversion has been known for a long time, but the interest toward this technology has not been constant for the whole period. Political and economic reasons as well as a more environmental friendly approach on energy production of the last decades have driven the world toward a greater attention to this topic. Early applications of the process were mainly for coal or fossil residues. Biomass gasification was practiced in the early decades of this century but received a major boost during World War II. Another wave of development began in the 1970s following the drastic increase in world oil prices and a growing awareness of possible climatic effects of continued use of fossil fuels. This new interest was accompanied by an expansion of the devices considered for use with biomass gasification. In particular, several applications, such as direct firing of gas in turbines and fuel cells, were expected to benefit from hot-gas cleanup of particulates and tar, at gasifier pressures.

One of the advantages of the gasification process is that pollutants can be easily removed from syngas before combustion. In this way, for example, the use of some high sulfur coals as the Sardinian Sulcis coal is much easier.

Typical use of syngas is energy production through an internal combustion unit or production of chemicals.

The main reactants in a gasification process are the fuel and the gasifying medium. Depending on the aim the gas feed may be constituted of air, pure oxygen and/or steam, sometimes a combination is also used. Main components of the syngas product are the permanent gases CO, CO2, H2 CH4 and vaporized tars; small amounts of unconverted char, soot and ash are also formed within the process. Soot, fly ash and tar may cause corrosion and blockage problems downstream the reactor; the development of an efficient gas cleaning system for particulate removal and tar conversion is, therefore, crucial in order to obtain a beneficial gasification process.

Gasification occurs in sequential steps: 1. heating and pyrolysis step where the fuel is converted into gas, char

INTRODUCTION

16

and primary tar; 2. cracking of primary tar; 3. heterogeneous gasification reactions of the char formed during

pyrolysis and homogeneous gas phase reactions; 4. combustion of char formed during pyrolysis and oxidation of

combustible gases. When a solid fuel is heated up to 300-600°C in the absence of an oxidizing

agent it decomposes to give char (a solid with a high carbon content), tar (condensable hydrocarbons), and gases. The relative yields of gas, tar and char depend mostly on the heating rate and the final temperature.

While the importance of this process is just pointed out as a step within a coal to energy route, pyrolysis has received increasing attention amongst the thermochemical biomass conversion routes. The solid fuel conversion to liquid depends on the pyrolysis process conditions: high heating rates, with rapid quenching favor the formation of liquid products and minimize char formation; high heating rates to high maximum temperatures favor the formation of gaseous products; slow heating rates coupled with low final maximum temperature increase the liquid and gas yields. The decreased formation of char at the higher rate of heating is accompanied by the increase on tar formation [9]. During pyrolysis fuel particles also undergo morphological and physical changes; this can also influence the overall devolatilization rate. Pyrolysis conditions also affect gasification reactions. It is well known that char reactivity for gasification and combustion reactions strictly depends either on the fuel type either on pyrolysis conditions.

Char produced from pyrolysis can further react with CO2, O2 and H2 to yield more gaseous fuels.

Generally pyrolysis proceeds much more rapidly than gasification. Pyrolysis products react with the oxidizing agent to give permanent gases (CO, CO2 and H2O) and lesser quantities of hydrocarbons. Char gasification is the interactive combination of several gas-solid and gas-gas reactions where solid carbon is oxidized to CO and CO2 while H2 is mainly generated through the water-gas shift reaction. The gas-solid reactions are the slowest and limit the overall rate of the gasification process, the more if a coal gasification process is attempted, it is indeed well known that the char produced from biomass is more reactive than the char derived from coal. Many of the reactions are catalyzed by the alkali metals present in ash.


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The gas composition is mainly influenced by the feed composition, water content, reaction temperature and the extent of oxidation of the pyrolysis products. Because of physical and geometrical limitations of the reactor and the chemical limitations of the reactions involved, tar produced during pyrolysis is not completely converted and contaminant tar is always present in the final product gas. These compounds tend to be refractory and are difficult to remove by thermal, catalytic or physical processes. This aspect of tar removal together with the reaction step is one of the most relevant technical issues of a gasification technology and is further discussed in Chapter 6.

2.2 An overview on the gasification process

A gasification system consists of three main sections: feed pretreatment, gasification, gas clean-up. A short description of every section is given in the following.

2.2.1 Feed pretreatment

Biomass is prepared in the forest usually as chips, although bundles from short-rotation forestry and whole logs from conventional forestry are delivered sometimes. Annual crops would be delivered in bundles (e.g. from miscanthus or sorghum), as bales (e.g. from straw), or possibly chopped (from any crop). This material has to be received, handled, stored and processed prior to gasification; the storing conditions are strictly related to the properties of the fuel at the feeding step (e.g. the moisture content).

Coal is commercially available in different particle sizes depending on the different destination. Among all the pretreatment steps milling is very important. Some technologies aiming at decreasing the undesired sulfur and inorganic matter are as well under development or have already been presented in literature but none of them has reached a strong potential for commercialization. These techniques are usually expensive and after coal processing a gas cleanup section is usually preferred.

2.2.2 Reaction step and reactor types

In order to convert the fuel, different gasifying agents are employed during

INTRODUCTION

18

the reaction step. One possible classification of gasification processes concerns the gas phase reagents [15]:

� Air gasification is the most widely used technology. A low heating value syngas containing up to 60% N2 having a typical heating value of 4–6 MJ/Nm3 is produced with this technology. This gas is suitable for boiler, engine and turbine operation but not for pipeline transportation due to its low energy density. The gasification temperature of an air blown gasifier usually ranges between 900-1100 °C;

� Oxygen gasification yields a better quality gas of heating value of 10–15MJ/Nm3 which is suitable for limited pipeline distribution and for further conversion (e.g. to methanol and gasoline). The difference between air and oxygen gasification is mainly due to the dilution effect of N2. In this process a temperature of 1000–1400 °C is usually achieved. Cost and safety problems are usually related to the requirement of an O2 supply.

In order to increase the hydrogen content of the system, for a better temperature regulation and the prevention of hot spots formation inside the reactor, steam is sometimes fed to the reactor, especially in the case of coal.

Since the use of air avoids costs and hazards related to oxygen production and usage gasification with air is the most widely used technology in the case of biomass.

Depending on reactor design, a further classification of the different technologies is often proposed. One possibility is to characterize reactor types on the type of transport of fluids or solids along the reactor. Three main classes are commercially available:

� quasi-non-moving or self-moving feedstock � mechanical-moved feedstock � fluidically-moved feedstock � other, usually constituted by a combination of the previous.

Many patented coal gasification technologies are commercially available; different new configurations for biomass gasification have been presented in the last years. Fixed bed and fluidized bed reactors are mostly used for gasification.

Fixed bed reactors can have countercurrent flow or co-current flow while crosscurrent mass flow is rarely used.


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In a countercurrent system the feedstock and the reactive material flow in opposite directions along the reactor while the opposite applies in regard to co-current. Depending on the direction of the reactive material it can be up- or down-stream: in the countercurrent reactor the gasifying agent is usually fed at the bottom and proceeds upwards (updraft configuration); in the co-current reactor the gasifying agent is usually fed at the top or middle of the reactor and the gas off take is in the bottom (downdraft configuration).

A downdraft gasifier features co-current flow between the gas phase and the solid phase through a descending packed bed supported across a constriction (throat), where the most part of gasification reactions take place. Reaction products are intimately mixed in the turbulent high-temperature region around the throat; there a well mixed region promotes tar cracking reactions. This configuration results in a high conversion of pyrolysis intermediates and hence a relatively clean gas. Downdraft gasification is simple, reliable and proven for a fuel with up to 30 wt% moisture content, an ash content less than 1 wt% and containing a low portion of fine and coarse particles (limiting dimensions range between 1cm and 30cm) [1]. Owing to the low tar content in the gas, this configuration is generally favored for small-scale electricity generation with an internal combustion engine.

In an updraft reactor the solid feed and the gas collection section are located at the top of the gasifier; ash discharge and gas feed are at the grate, near to the bottom; the flow of the two phases is countercurrent. During its way along the reactor the fuel is first dried and then decomposes to give tar, char and gas by means of pyrolysis reactions. Then, char and gas produced further react by means of gasification and combustion reactions yielding more gas. Since tar, once formed, moves towards lower temperature sections it does not decompose before leaving the reactor. Updraft gasification is the gasification technology with the higher tar concentration in the product gas and the needier of an efficient cleaning system. At the bottom of the gasification zone the solid char from pyrolysis and tar cracking is partially oxidized by the incoming air or oxygen. The significant proportion of tars and hydrocarbons in the product gas contributes to its high heating value. The principal advantages of updraft gasifiers are their simple construction and high thermal efficiency: the sensible heat of the gas produced is recovered by direct heat exchange with the entering feed, which is thus dried, preheated and pyrolysed before entering the

INTRODUCTION

20

gasification zone.

Fluidized bed gasification is usually carried out with bubbling bed or circulating bed and entrained flow technology.

Inside fluid bed gasifiers the advantage of the excellent mixing properties and high reaction rates of this gas-solid contacting technology is exploited. The fluidizing material is usually silica sand, although alumina and other refractory oxides have been used to avoid sintering. Different catalysts have also been tested inside the reactor in order to reduce tars during the reaction step. A typical operating temperature for biomass gasification ranges between 800-900°C. The reactor is in this case constituted by a bed bounded at the bottom with a gas distribution plate and a freeboard section. Feedstock can be fed at the top or at the bottom of the reactor. Most of the conversion takes place within the bed; hence some conversion to product gas continues in the freeboard section especially reactions of tar cracking. Sometimes an excessive carryover of fines takes place, especially in a top feeding configuration. When biomass is fed to the reactor the bubbling fluid bed gasifier tends to produce a gas with tar content between that of the updraft and that of the downdraft gasifier. Depending on thermal characteristics of ash the bed and the distribution plate sintering is one of the commonly encountered problems. Alkali metal compounds from biomass ash form low-melting eutectics with the silica in the sand resulting in agglomeration and bed sintering with eventual loss of fluidization. However, the lower operating temperature of a fluid bed and its better temperature control related to the isothermal operation provides suitable environment for the gasification of many biomass materials. Fluidized bed gasifiers can also be scaled up with considerable confidence. Alternative configurations such as circulating fluidized beds and twin-bed systems are also available. In the circulating fluid bed is produced a hot raw gas which is very often used for close-coupled process heat or in boilers to recover the sensible heat in the gas. This configuration has been extensively developed in the case of biomass and many commercialized applications for coal processing also exist.

A twin fluid bed gasifier is sometimes used in order to obtain a gas of higher heating value than that obtained from a single air-blown gasifier. The gasifier is thus constituted by two reactors: a pyrolyser and a char burning reactor, the first is heated with hot sand from the second fluid bed. Product quality is good as it has a high heating-value but has a high tar loading because the pyrolysis process


21

takes place at low temperature.

In entrained flow gasifiers no inert material is present but a finely reduced feedstock is required. Entrained bed gasifiers operate at quite higher temperatures (1200-1500°C) the thermal cracking of tar is favored in this condition and low concentrations of tars and condensable gases exist in the product gas. Nevertheless, this high-temperature environment, very close to the ash melting point, originates sintering problems and difficulties with the choice of the constituent material of the reactor.

Each type of reactor has its advantages and limitations. The main advantages of fixed bed reactors are related to the high carbon conversion efficiency, the wide range of ash content in the feedstock and the possibility to melt ash; this is indeed discharged as dry ash or as a molten slag depending on the technology adopted. At the same time they have some start up problems and difficulties on the scale up of the reactor. Nevertheless, fluidized beds have: a better heat and material transfer between the gas phase and solid phase, a good temperature distribution, high specific capacity and faster heat-up. Fluidized beds are still unfavorable for their high dust content in the gas phase and the conflict between high reaction temperatures needed for a good conversion efficiency and low melting points of the ash components.

In fixed bed gasifiers with IC engine applications high concentrations of particulates and tars can damage the engine or originate a too high level of maintenance. Nevertheless, small-scale atmospheric fixed bed gasifiers are usually selected as gas producers for IC engine power generators [1].

The amount of tar is much higher in countercurrent than in concurrent gasifiers, countercurrent gasifiers are therefore not considered as an option for internal combustion engine applications.

2.2.3 Gas clean up section

Gases formed during gasification are contaminated by many pollutants. The contamination level varies depending on process type and feedstock. Gas cleaning must be applied in order to prevent erosion, corrosion, blockage, catalyst poisoning and environmental problems in downstream equipment.

Gas streams contain very small particles that are difficult to remove by

INTRODUCTION

22

cyclones, a scrubbing section is sometimes suitable, but hot gas cleanup is for example important for pressurized systems, where the sensible heat of the gas needs to be retained and scrubbing systems for tar removal avoided. High-temperature ceramic or metal candle filters have been tested with gasification products from peat and coal and are commonly used within biomass gasification processes.

One of the most important problems of biomass gasification is the tar content of the gas product. Tar concentration is mainly a function of gasification temperature and its level is therefore a function of reactor type and processing conditions. Tar formed during pyrolysis is thermally cracked in most environments to refractory tars, soot and gases. Tar level and characteristics are also slightly dependent on the feedstock. The most important problem connected to the tar content of the fuel is related to condensation and reaction to give soot which can block filters. Tests have shown that tar production in wood gasification is much greater than in coal or peat gasification and it is also demonstrated that even the tar composition differs between the case [1, 2]. This implies that technology developed in coal gasification tar cleaning may not be directly transferable to biomass feeds and a lot of work has been done in order to understand the behavior of biomass tar and to develop an effective and selective tar removal section. Problems related to the tar content of the gas are discussed in Chapter 6.

Alkali metals are the major ash components of many biomass forms although the particular nature of biomass-derived alkali metals and their association with other contaminants such as sulfur is not known. These compounds, at high temperature, exist in the vapor phase and are therefore able to pass through particulate removal devices. Alkali metals cause high-temperature corrosion of turbine blades, they can indeed strip off their protective oxide layer. Alkali metals may also damage ceramic filters at high temperature and the gas will thus need to be cooled off before the filter.

During gasification the most part of fuel-bound nitrogen is converted to ammonia and smaller quantities of other gaseous nitrogen compounds causing NOx emissions. Since the nitrogen-containing contaminants are in the vapor phase the most part of particulate removal devices are useless for them. The reduction of NO production is possible by limiting fuel-bound nitrogen in the


23

feedstock through the selection of biomass type, wet scrubbing of the product gas and, whenever possible, without a large heat loss of sensible heat stored in the gas phase. Some low-NO, combustion techniques are also used in order to minimize thermal NO production and selective catalytic reduction is adopted whenever a high NOx content is present at the exhaust of the engine or turbine.

Sulfur is not generally considered to be a problem for biomass gasification since biomass has very low sulfur contents. Sulfur is normally not associated to biomass but what is often viewed or claimed as ‘trace’ level of, for example, 0.1 wt%, can lead to sulfur concentration in the gas up to 100 ppm. This level is not acceptable and requires reduction since the specification for turbines is typically 1 ppm or often much less especially when the gas is used for chemicals production where the most part of catalysts are easily poisoned by sulfur [1].

Expensive sulfur removal trains are not necessary for biomass gasification and especially if a dolomite tar cracker is included in the process this is significantly absorbed. Nevertheless, sometimes the low levels required are not achieved and a sulfur monitoring is always necessary. A sulfur guard, consisting of a hot fixed bed of zinc oxide, is likely to be adequate for the concentrations expected. This is relatively inexpensive to install but creates a waste disposal problem from the zinc sulfide produced or further costs for zinc sulfide regeneration.

Since a large part of dirty coals exist the problem is more relevant in case of coal gasification. For example the Sulcis coal, produced in the South-East of Sardinia, has a high sulfur content of about 6-7% [5]. Sulfur removal may therefore be necessary. Two routes can be attempted: hot or cold gas cleaning. Desulphurization processes commercially available are usually based on liquid scrubbing at or below room temperature. Developments in hot-gas desulphurization have focused on re-generable solid metal oxides as promising candidates sorbents. The importance of hot gas cleanup is increasing its importance because of the opportunity to keep the heat content of the gas in order to do not decrease the thermal efficiency of the whole process.

Chlorine and its compounds are also present in the gas; they can be removed by absorption in an active material either in the gasifier or in a secondary reactor, or by dissolution in a wet scrubbing system. Dolomite and related materials are less effective at removing chlorine than sulfur. Chlorine is another potential contaminant in case of biomass; it can arise from pesticides, herbicides

INTRODUCTION

24

and in waste materials.

2.2.4 Influence of parameters

Unlike coal, where high pressure offers considerable performance advantages, pressurized systems have a lower volume but there is relatively little reaction-kinetic or thermodynamic advantage resulting from pressurized operation with biomass. It is indeed so reactive that these advantages do not compensate the increase in plant size and fixed cost related to investment.

Feeding for pressurized biomass gasification is more complex, very costly and has a high inert gas requirement for purging. Capital costs of pressurized equipment are much higher than for atmospheric equipment, although equipment sizes are much smaller. This disadvantage is countered by the higher efficiency. As the gas is supplied to the turbine at high pressure the need for gas compression is removed. Relatively high tar content in the gas is also permitted for pressurized systems.

Atmospheric systems have a potentially much lower capital cost at smaller capacities while gas compositions and heating values are not significantly different for either system.

Since its connection to the energy amount used to run the reactor the oxygen to fuel ratio is the most important parameter for coal and biomass gasification. This parameter is usually compared to the stoichiometric oxygen requirement for total carbon burning: the ratio between the oxygen fed to the reactor and the oxygen requirement for total combustion is very often named equivalent ratio and indicated as λ. When λ is equal to zero the system is allothermal and the heat required for reaction is provided with different methods such as heated walls or electrical heaters, for an equivalent ratio equal to one or more the system is a combustor while intermediate values are typical of gasification. As a heuristic it is said that equivalent ratio in the case of gasification ranges between 0.25 and 0.35. This is an absolute value as it is referred to the stoichiometric oxygen; the oxygen to fuel ratio, or the air to fuel ratio in the case of air blown gasifiers are the real operating parameters.

The steam to fuel ratio is also important. A better temperature control inside the bed and the promotion of gasification reaction is expected with an increase of this parameter. Sometimes the steam flow is referred to the oxygen /air flow,


25

a steam to air ratio is in that case defined.

Since the process is intrinsically heterogeneous the particle affects the process. Smaller is the particle size better is the heat transfer. Temperature is uniform resulting in reaction taking place throughout the particle. Because of the anisotropic nature of wood and its electrostatic properties particles shape plays an important role for biomass, one common encountered problem is feeding in indeed the fuel feeding section. Nevertheless pelletization and chipping consume power decreasing the available energy. A proper analysis is required for the particular case in order to find the best choice.

Porosity affects gasification reactions, if the fuel has a high void fraction the surface area for reaction is very high and the diffusion of the reactant/ product would be easy. Uniform temperature could be achieved throughout biomass resulting in continuous reaction at all portions of biomass yielding uniform composition of product gases. When biomass is less porous, temperature varies from the maximum at the exterior to the minimum at the interior; reaction takes place only at the exterior surface that shrinks with reaction. Because of the non-uniformity in temperature drying, pyrolysis and gasification take place simultaneously yielding non-uniform gas composition.

Since pyrolysis and gasification of biomass are thermochemical processes, temperature and heating rates have pronounced effects on the weight loss of biomass. In fast pyrolysis, with high heating rates of up to 1000 °C/min at temperature below 650 °C and with rapid quenching the liquid, intermediate products of pyrolysis condense without further breaking down higher molecular weight species into gaseous products. Formation of char is minimized by high heating rates. Gaseous products form at high heating rates and high maximum temperature. If the desired end product is the liquid hydrocarbon or bio-oil, fast pyrolysis is preferred [8]. Fast pyrolysis is typical of fluidized bed, cyclonic, entrained flow, vortex and ablative reactors. Slow pyrolysis, typical for fixed bed reactors, requires low heating rates and low maximum temperature. Maximum char yield via secondary coking and repolymerization reaction is observed at slow heating rates coupled with a low final maximum temperature and with long gas and solid residence times. Slow pyrolysis is commonly used for char production.

INTRODUCTION

26

The ash content of the fuel is also an important parameter. Many works report its influence on gasification reactions. It is well known that the presence of even very low concentrations of impurity either catalyzes or inhibits the degradation [8,16].

Cellulose, hemicelluloses, lignin and extractives are found to be the major components of biomass. The thermal degradation of biomass and samples of cellulose, hemicellulose and lignin has been studied extensively using TG analyzer.

From a chemical point of view both of them, biomass and coal are constituted by carbon, hydrogen and oxygen in major quantities. The main difference between the two fuels, with reference to the chemical composition, is related to the oxygen content. Each biomass has oxygen to convert to carbon oxides either partially or completely, this causes a higher volatiles production and therefore a higher importance of the devolatilization step in case of biomass processing.

2.3 References

1. Bridgwater A. V., The technical and economic feasibility of biomass gasification for power generation, Fuel, 74, 1995, 631-653.

2. Brage C., Yu Q., Chen G., Sjöström K., Tar evolution profiles obtained from gasification of biomass and coal, Biomass and Bioenergy, 18, 2000, 87-91.

3. Collot A.G., Matching gasification technologies to coal properties, International Journal of Coal Geology, 65, 2006, 191-212.

4. Demirbaş A., Biomass resource facilities and biomass conversion processing for fuel and chemicals, Energy Conversion and Management, 42, 2001, 1357-1378.

5. Fois E., Pistis A., Melis F., Pisanu F., Deriu G., Mura G., The removal of mineral matter and sulphur from Sulcis coal by leaching, Proceedings of the 4th International conference on Clean Coal Technologies, Dresden, 2009.

6. Hasler P., Nussbaumer Th., Gas cleaning for IC engine applications from fixed bed biomass gasification, Biomass and Bioenergy, 16, 1999, 385-395.

7. Hobbs M. L., Radulovic P. T., Smoot L. D., Combustion and gasification


27

of coals in fixed-beds, Progress in Energy and Combustion Science, 19, 1993, 505-586.

8. Kirubakaran V., Sivaramakrishnan V., Nalini R., Sekar T., Premalatha M., A review on gasification of biomass, Renewable and Sustainable Energy Reviews, 13, 2009, 179-186.

9. Küçük M. M., Demirbaş A., Biomass conversion process, Energy Conversion and Management, 38, 1997, 151-165.

10. Longwell J. P., Rubin E. S., Wilson J., Coal: energy for the future, Progress in Energy Combustion Science, 21, 1995, 269-360.

11. Mills J. S., Coal gasification and IGCC in Europe. IEA Coal Research (CCC/113), 2006.

12. Milne, T.A., Abatzoglou, N.,Evans, R.J., Biomass Gasifier “Tars”: their nature, formation and conversion. Report NREL/TP-570-25357, NREL, Golden, Colorado, USA, 1998.

13. Moreea- Taha R., Modelling and simulation for coal gasification, IEA Coal Research (CCC/24), 2000.

14. Qader S. A., Coal science and technology - Natural gas substitutes from coal and oil, Amsterdam, 1985, Enselvier.

15. Saxena R. C., Seal D., Kumar S., Goyal H. B., Thermochemical routes for hydrogen rich gas from biomass: a review, Renewable and Sustainable Energy Reviews, 12, 2008, 1909-1927.

16. Scott D. H., Ash behaviour during combustion and gasification, IEA Coal Research (CCC/24), 1999.

17. Smith I. M., Rousaki K., Prospects for co-utilisation of coal with other fuels – GHG emissions reduction, IEA Coal Research (CCC/60), 2002.

18. Warnecke R., Gasification of biomass: comparison of fixed bed and fluidized bed gasifier, Biomass and Bioenergy, 18, 2000, 489-497.

19. Wender I., Reactions of synthesis gas, Fuel Processing Technology, 48, 1996, 189-297.

20. Wu Z., Fundamentals of pulverized coal combustion. IEA Coal Research (CCC/95), 2005.

Section I

A mathematical model for a fixed bed

updraft reactor

A MATHEMATICAL MODEL FOR A FIXED BED UPDRAFT REACTOR

31

Nomenclature

∆H av C Cp h K k L R

Rcg S T U z

Greek letters

ρ ε ζ

Superscripts

max * ni

Subscripts

c cg d g gc sw gw m

heat of reaction reactor specific area concentration heat capacities heat transfer coefficient equilibrium constant kinetic constant bed height kinetic rate ideal gas reactor cross area Temperature heat transfer coefficient reactor abscissa

density gas phase solid phase

void fraction multiplicative coefficient for the gas solid heat exchange

maximum yeld real reaction order

coal char-gas reaction drying gas gas phase combustion solid – wall gas – wall

J/Kg m2surface/m3reactor

mol/m3gas J/(Kg K) W/(m K

- - m -

J/(mol K) m2

W/(m K) K m

Kg/m3gas

Kg/m3reactor m3void/m3reactor


32

wp cp r s T t w wg

moisture wood pyrolysis coal pyrolysis reaction solid tar time wood water gas shift


33

Introduction

The reactor is the heart of every chemical process since its performance influences the whole system. A deep understanding of its behavior is important in order to obtain the desired products at different operating conditions. This aim can be achieved with an experimental investigation or with a modeling approach, the interaction of both should be preferred.

Mathematical modeling is a powerful tool either for process design either for operating equipment. It can be used to simulate its performance and it is very helpful in order to understand the dynamic of the system and design the best control system for the process.

The development of a model of a countercurrent fixed bed gasifier for a deeper understanding of the transient behavior of the reactor, the study of the steady state condition and the investigation of the influence of some operating parameters is the aim of this work. Two cases are considered: coal feed and biomass feed to the reactor.

A number of models for gasification systems are available in the literature. Most of them refer to fluidized bed reactors but a number of works dealing with fixed bed description have been also presented.

Although the most part of existing gasifiers, especially for coal conversion, are fixed bed reactors modeling of this system has not received particular attention over the years. Previous works about a coal fed countercurrent gasifier are present in the literature [6, 7, 9, 12-14]. The most part of them is one dimensional but even some two-dimensional model exist. The process is studied from the equilibrium point of view [11] or with a kinetic approach [1-12, 14] were:

� single shape and size of the particles; � no momentum transfer; � constant porosity of the bed; � heat and mass transfer coefficients for non reacting systems; � instantaneous drying; � instantaneous or highly simplified solid devolatilization; � uncertainty on the intrinsic kinetics of heterogeneous combustion

and gasification reactions;


34

� no homogeneous gas-phase reaction, apart from the water-gas shift equilibrium;

� steady, one-dimensional equations; � limited model sensitivity analysis and validation.

They also neglect the influence of the ash content. Ash is not even mentioned in some works.

A rather good and detailed model for coal conversion is presented in [6]. Authors published a first model and then the improvement in a further work [9].

A comprehensive review on coal fixed bed modeling up to 1993 was presented by the same authors [7]. After that work the interest toward this system decreased and studies dealing with this topic have almost disappeared.

Few models of biomass fed updraft reactors exist in the literature [3, 8, 12]. A model dealing with the transient behavior of a Lurgi type gasifier exist [14]. Almost all the models of fixed bed gasification found in the literature

describe the steady state behavior of the reactor. Some previous works tried to attempt a description of the unsteady state of biomass gasifier [3, 12].

In particular in [3] is described the most comprehensive model for this system. It still does not consider the possibility to feed steam. The dynamic behavior of the bed composition is neglected and the sensitivity analysis is conducted in a small range of operating parameters. The dynamic behavior is described only with regard to the combustion zone and just for a short time; the work does not report any other information about the transient behavior.

A work dealing with high temperature gasification in a fixed bed is also available in the literature [12], in this work the problem of the ignition of the bed is analyzed, but once more the behavior of the bed, especially for long time is neglected.

The work presented in this paper is a pseudo homogeneous one dimensional model of an updraft reactor with a kinetic approach in the description of the different phenomena involved. It is composed by a set of differential and algebraic equations derived from mass and energy continuity principles. A study of the dynamic behavior of the reactor is here presented; moreover, the changes within the bed and the influence of the operating parameters are studied in a wide range. See Chapter 3 for a detailed description of the model.


35

References

1. Amundson N. R., Arri L. E., Char Gasification in a countercurrent reactor, AIChE Journal, 24, 1978, 87-101.

2. Caram H. S., Fuentes C., Simplified model for a countercurrent char gasifier. Industrial and Enginering Chemistry Fundamentals, 21, 1982, 464-472.

3. Di Blasi C., Modeling wood gasification in a countercurrent Fixed-Bed Reactor, AIChE Journal, 50, 2004, 2306-2319.

4. Gao, N., Li, A., Modeling and simulation of combined pyrolysis and reduction

zone for downdraft biomass gasifier. Energy conversion and management, 2008, 49, 3483-3490.

5. Gøbel, B, Henriksen,U., Jensen, T.K., Qvale,B., Houbak, N.,(2007). The development of a computer model for a fixed bed gasifier and its use for

optimization and control. Bioresource Tecnology, 98, 2007, 2043-2052. 6. Hobbs M. L., Radulovic P. T., Modeling fixed-bed coal gasifiers, AIChE

Journal, 38, 1992, 681-702. 7. Hobbs M. L., Radulovic P. T., Smoot L. D., Combustion and gasification

of coals in fixed-beds, Progress in Energy and Combustion Science, 19, 1993, 505-586.

8. Mandl C., Obernberger I., Biedermann F., Updraft Fixed-bed gasification of softwood pellets: mathematical modeling and comparison with experimental

data, Proceedings of the 17th European biomass conference and exhibition, Hamburg, 2009.

9. Radulovic P.T., Ghani M. U., Smoot L. D., An improved model for fixed bed coal combustion and gasification, Fuel, 74, 1995, 582-594.

10. Tinaut, F. V., Melgar, A., Pérez,J. F., Horrillo, A., Effect of biomass particle size and air superficial velocity on the gasification process in a downdraft

fixed bed gasifier. An experimental and modeling study. Fuel processing technology , 89, 2008, 1076-1089.

11. Tola, V., Cau, G., Process analysis and performance evaluation of updraft coal gasifiers. Third International Conference on Clean Coal Technologies for Our Future, Cagliari, Sardinia, Italy , 15-17 May 2007.

12. Yang W., Ponzio A.,Lucas C., Blasiak W., Performance Analysis of a fixed-bed gasifier using high temperature air, Fuel Processing Technology, 87, 2006, 235-245.


36

13. Yoon H., Wei J., Denn M. M., A model for moving-bed coal gasification reactors, AIChE Journal, 24, 1978, 885-903.

14. Yoon H., Wei J., Denn M. M., Transient behavior of moving-bed coal gasification reactors, AIChE Journal, 25, 1979, 429-439.

Part

I

The model

Chapter

3

Model description

A lack of studies dealing with countercurrent fixed bed reactors exist in the literature, especially with reference to the dynamic behavior of the gasifier. This fact led to the need to fill up the lack and attempt the description of the system.

The one dimensional phenomenological model developed during a three year study of a fixed bed countercurrent gasifier is described in this chapter. A kinetic approach is used to describe phenomena involved in the process. Heterogeneity of the system is somehow counted in while it is in some case neglected. An exhaustive description of assumptions, constitutive equations and all the necessary information for the understanding of the model are presented in this Chapter.

CHAPTER 3 MODEL DESCRIPTION

41

3.1 Introduction

A one dimensional dynamic model of an updraft rector based on a set of differential equations derived from heat and mass continuity principles is here described. Two different approaches are commonly used to address the description of a fixed bed gasification system: equilibrium or a kinetic approach involving the reaction rates of the different phenomena considered. A mixture of both methods is sometimes used; in that case a kinetic description of pyrolysis and gasification reactions is usually attempted while the kinetic equilibrium is often considered for those reactions that take place in the gas phase. A kinetic approach is used in this work.

3.2 Governing equations

A mass balance for each component involved in the process either in the gas or in the solid phase is considered. Differently from many works previously presented in the literature two separate heat balances, one for the gas and one for the solid phase is used. A detailed description of the assumptions that constitute the base of the model is given below while a summary of governing equations is shown in Table 1 and Table 2.

3.2.1 Mass balance

Biomass tar, coal tar, CO, CO2, H2, CH4, H2O, O2, N2 are the gas species counted in the model. According to the EU/IEA/US-Doe meeting held in Brussels in 1998, dealing with the definition of a tar sampling protocol, tar is defined as all organic contaminants with a molecular weight larger than benzene [34].

However, most part of kinetic models for pyrolysis presented in the literature assume CO, CO2, H2, CH4, as the gas products of pyrolysis and heavier hydrocarbons are referred as tar. In order to allow the use of global kinetic models, in this work, tar is considered as a mixture of hydrocarbons with a molecular weight higher than CH4.

Two main differences exist between the tar produced from coal and that produced from biomass: the first deals with the amount, larger quantities of

A MATHEMATICAL MODEL FOR A FIXED BED UPDRAFT REACTOR THE MODEL

42

condensables are produced from biomass than from coal; the second is related to its composition, heavier aromatic compounds are usually present in coal tar. For further description of tar composition see Chapter 6.

Although the difference existing between tar yielded from biomass and that yielded from coal, tar is assumed as an aromatic compound with formula

��.��.� and a molecular weight equal to 95 ��⁄ [31]. Naphthalene is chosen as the reference compound for thermal and transport properties. Although this is more likely for biomass than for coal this assumption is not expected to influence.

Depending on type of fuel, storage conditions and feed pre-treatment, moisture in form of interstitial and bound water, is usually present. Bound water is contained in both, coal and biomass, interstitial water is more likely to be present just in biomass. Steam is released from the fuel if exposed to high temperature environment and its presence affects gasification, especially gas phase reactions. Since evaporation is an endothermic process moisture release also affects the bed temperature.

In this work the moisture content of the fuel is considered as a fraction of the solid phase.

The ash content of the fuel usually refers to all those compounds that do not participate on gasification and combustion reactions or, in detail, to the residue remaining after the complete combustion of the fuel. Despite of their change in structure, reaction with other compounds (e.g. Sulphur) and their catalytic properties, they can be referred as inert in a gasification process. A number of models dealing with the description of a fixed bed reactor behavior neglect ash influence, they do not even mention the presence of ash within the bed. Nevertheless, as it influences gas solid reactions, its description is very important for a deep understanding of the bed behavior and its changes during reaction. They have a catalytic action toward gasification and influence the specific area. For this reason, in order to investigate the role of ash inside the packed bed, a reacting and an inert fraction of the fuel are considered and respectively named as wood (or coal depending on the feed) and ash within the model.

The reacting content of the fuel is calculated as the sum of fixed carbon and volatile matter while ash refers to the ash content of the fuel as indicated in the proximate analysis.


43

As discussed before, during pyrolysis the solid fuel decomposes to give a solid phase and a gas phase fraction. The solid phase is commonly referred as char. Char is a highly porous solid material with a high carbon fraction. It is very difficult to find an univocal definition or univocal properties for char because it depends on its formation process. The original fuel is the first factor affecting its properties: it is well known that char prepared from coal is less reactive than this yielded from biomass pyrolysis [10, 29, 56]. Even when the same fuel is used a different char is obtained at different pyrolysis conditions. Although many studies on char reactivity and qualitative information about char properties are available in the literature, it is still difficult to insert it in a model. Biomass char and coal char are considered to be the same compound within this model.

After pyrolysis the solid still contains small amounts of hydrogen and oxygen, nevertheless char is considered as pure carbon during calculations

The description of the main phenomena involved in the process have been considered in this study and a first attempt to a dynamic model of an updraft reactor is the aim of the work; further improvements (e.g. the description of the fate of sulfur and nitrogen) are left to descendants. Solid compounds considered by the model are: moisture, wood/coal, ash and char.

3.2.2 Heat balance

Unlike the case of a fluidized bed reactor, where homogeneous thermal conditions inside the bed can be assumed, thermal conditions along an updraft reactor may vary significantly. A hot zone is located very close to the grate while temperature is expected to decrease along the reactor height until the biomass feeding section.

The most part of phenomena involved in the process is thermally driven. It is thus well understood that temperature plays an important role during gasification, consequently a comprehensive description of the reactor behavior cannot abstract from a good description of the thermal behavior of the two phases along the bed. Since an air feed is provided at the grate, combustion reactions are promoted in this zone (or where the fuel comes in contact with oxygen) and the process is autothermic. Heat is produced in a small section while it is required along the whole bed. In the upper part of the gasifier the heat supply to the solid phase is closely related to the ability of the gas phase to work as a heat carrier along the bed: a good description of the heat exchanged


44

between the two phases is required. Two separate heat balances are considered: one for the gas phase and one for the solid phase. A convective and radiation heat exchange term between the gas and the solid phase is considered and

calculated according to [4]. A dimensionless correction factor � is also introduced to account for uncertainties related to the presence of a reacting bed. The heat transfer mechanism assumed for the solid and the gas phase is conduction.

The system considered in this paper is equipped with a cooling jacket fed with water at atmospheric pressure. For modeling purposes a fixed temperature of the gasifier wall equal to 100°C is assumed and heat losses through the wall are calculated by the model [18].

3.2.3 Other assumptions

Stationary solid or plug flow along the bed is the most common assumption in the description of the solid flow along the bed. Some sub-models assimilating the solid movement to the flow of a granular material exist but the topic still needs more investigation and at the state of the art they would be too complex. If used in a comprehensive reactor model, they would influence dramatically the calculation time [24].

Since a relation between the gas flow, bed porosity and the different conditions along the reactor exist, the plug flow assumption is inadequate for this system, channeling was also affects the flow in fixed bed reactors [24].

Constant velocities within the section is assumed for both, the gas and the solid phase but variable axial velocities are considered and respectively calculated through the overall mass balance for each phase in order to take into account the changes in physical conditions, composition and porosity along the reactor.

According to the definition of bed porosity:

� = 1 − � ��∗

��

��

In the last equation � is the bed porosity, �� is the apparent density of the

component and ��∗ is the real density of the material; the bed density is calculated as a function of bed composition.

Ideal gas behaviour was assumed and ideal gas state equation is used to


45

calculate the total gas density at each section. Solid and gas physical properties and particle dimension are calculated in

every section, a variation with temperature along the reactor according to proper relations available in the literature is considered.

For the ash behavior description two main sub-models for particle radius calculation are available in the literature [24]. According to the ash segregation model (AS), ash crumbles and falls away, in this case the oxidant is required to diffuse through the boundary layer. According to the shell progressive model (SP), ash remains intact and oxidant is required to diffuse through the boundary layer and through the ash layer. In the AS model ash is assumed to segregate and form ash particles with fixed diameters, then a constant or variable particle number density can be assumed. The first assumption is likely for stationary bed but not for a moving bed. The SP model assumes constant or variable particle size and number density. The SP model with variable particle size as a function of conversion and variable particle size was assumed in this model. Atmospheric pressure is assumed inside the gasifier and pressure variation along the bed is not included in the model.

Nevertheless, equilibrium considered for water gas shift reaction (see

Table 1: constituting equations: mass balances.

MASS BALANCES

Solid phase component

− �� ! + � #�,%

�&

%��= ��

�' ( = 1, … , *

Overall solid phase

− � �� !

��

��+ � � #�,%

�&

%��

��

��= � ��

�'��

�� ( = 1, … , *

Gas phase

−+,��!� + ��

�! + � #�,%

�&

%��= ��

�' ( = 1, … , *-

Overall gas phase

− � +,��!�

�.

��+ � ��

�!

�.

��+ � � #�,%

�&

%��

�.

��= � ��

�'

�.

�� ( = 1, … , *-


46

Table 2: constituting equations: heat balances.

paragraph 3.3.3) should not be affected by the pressure drop since the reaction occurs with a constant number of moles, all the other reactions are assumed to be at kinetic or diffusive control and a small pressure variation is not likely to significantly influence the process.

The increase of the void fraction due to the disappearance of the solid phase through pyrolysis and gasification reactions should be considered. A lack of information on bed variation along the gasifier is present in the literature, therefore scarce data is available about void fraction within the bed. A void fraction equal to 0.3 at the top of the reactor and 0.7 at the bottom was indicated in [24]. The same authors but in another work dealing with the gasification of a number of U.S. coals reported just a little variation in porosity between solid inlet and outlet [25].This phenomenon can be explained by the competition between the increasing void fraction due to the conversion of the solid and the compression of the bed by the upper solid weight. The reactor shape, operating pressure and temperature influence bed porosity. In particular at high temperature the gas produced inside the pores during gasification reactions can expand and increase porosity. Depending on the weight of the bed particles can crash or keep the new void fraction. This phenomenon is very common for fluidized bed reactors where single particles are not subjected to the weight of the bed above but just to impacts to the wall and to other particles that could also lead to particles crash. Most part of the previous models found in the literature consider the increase of void fraction with a linear variation

HEAT BALANCES

− � �� /01��2 �!

��

��+ � ∆�%#%

�&�

%��− 456782 − 2-9

− 4+ ℎ <=2 − 2<> − ? ��2

�!� = � �/01��2 �'

��

��

Solid phase

� ��-/01��2-�!

�.

��+ � ∆�%#%

�&.

%��+ 456782 − 2-9

− 4+ ℎ <82- − 2<9 − ? ��2-

�!� = � �/01��2-�'

�.

��

Gas phase


47

along the reactor height, the assumption of a constant void fraction along the bed is as well diffused. A constant void fraction is assumed in this work.

An air blown gasifier is here considered, the possibility of steam injection to the reactor is also inserted. In this case a perfect mixing of the two streams is assumed.

3.3 Kinetic description

Reactants of a gasification system are in the solid and in the gas phase. For this reason the process is intrinsically heterogeneous. A number of phenomena should be included in the description of gas-solid reactions: reactant diffusion inside the bed and sorption inside the particle. Nevertheless, for sake of simplicity and in order to avoid a long lasting computation time, a complete description of the heterogeneousness of the system is far from being suitable in this work and a pseudo homogeneous approach is assumed for the kinetic description of the system.

3.3.1 Drying

Depending on storage conditions and fuel pre-treatment moisture is always present inside biomass. Drying is the first step of a gasification system and therefore a description of the process has to be inserted in a comprehensive reactor modeling. Although a detailed description of moisture loss and steam diffusion inside the particle would be appreciable, it is outside of our interest and cannot add any further improvement to the overall work. Nevertheless a poor description of the process with an instantaneous release is far from being real so a method that considers the dependence of drying on moisture content and environmental conditions has to be employed. In accordance with previous studies dealing with fixed bed modeling the process is rounded with a first order kinetic law depending on moisture content of the bed [9].

In the case of big particles non uniform thermal conditions inside the particle exist and the steam already yielded from moisture evaporation can thus condense if exposed to lower temperature. However this phenomenon has to be considered in the description of the gasification behavior of a solid particle, where non uniform thermal conditions are studied, while it is negligible or asserts a lower influence when a constant temperature along the whole section is


48

considered.

3.3.2 Pyrolisys

A solid fuel, once exposed to a high temperature environment decomposes to give char tar and gas. In general the process is widely recognized to occur in at least two main steps: a first solid decomposition to give char, tar and gas, commonly referred as primary pyrolysis, and secondary pyrolysis where tar further reacts yielding more gas and char.

Due to the number of phenomena involved in pyrolysis an investigation of this process is very difficult and a detailed description of the reacting system and its behavior is therefore an ambitious goal. Distribution between the different products during devolatilization reactions may vary depending on many factors. Among them the heating rate, the particle dimension and the kind of solid employed are the most important.

A detailed and exhaustive description of this process is very hard to achieve and far from our purpose, we are indeed looking for a compromise between a good kinetic description and a short computational time. A brief description of the kinetic models available in the literature for both, biomass and coal is presented in the following and the model chosen for this work is indicated.

3.3.2.1 Coal

As the result of years of investigation hundreds of models of coal pyrolysis are available in the literature. They can be mainly classified in two big fields: very detailed models comprehensive of many information (e.g. gas composition, char production and sometimes even some structural properties of char) and more simple global models. FLASHCHAIN [38], CPD [21] and FG-DVC [57] belong to the first category. These models are quite comprehensive: they describe devolatilization as a depolymerization into smaller volatile fragments and further reaction to give char. Differently from the first two FG-DVC counts in secondary pyrolysis reactions as well. All these models include a set of sub-models for the description of coal composition and structure. Though these models can predict yields, transient evolution rates and other information for a wide range of operating conditions, they need a large number of parameters in order to be used (e.g. CPD and FG-DVC also need C NMR analyses). Another


49

problem related to these models is the need to use the software properly prepared from the authors and the difficulty to match their output with a comprehensive gasification model.

On the other hand the application of global kinetic models is very simple but at the same time they are not able to provide as much information as the previously described network models. They are usually specialized on a singular coal and particular heating conditions and it is very hard to find a work that really suites for many different systems or that allows some generalization [1, 3, 23, 33, 49, 52, 53]. Moreover, results presented in such works are very often incomplete, the behavior of the gas is sometimes described while the solid phase neglected or vice versa.

These models can be divided into three main classes [52]: simple one step models, competitive reaction models (CRM) and multiple reaction models (MRM). An Arrhenius kind expression for temperature dependence is always considered while differentiation exists between the models with relation to the products or reactants dependence. The first model usually revolves on coal conversion, the reaction rate is thus calculated through the amount of residual coal.

The second model considers the competition of multiple reactions, each one yielding a different gas species. The reaction rate is in this case calculated with a first order kinetic law depending on the single component yield in the gas phase and the maximum yield of the process for this kind of coal.

Within the third model a continuous distribution of the activation energies of the individual reactions (DAEM) model is considered.

Because of the way they are formulated the use of global models leads to the problem of stoichiometric coefficients determination and can originate some problems in the elements balances during reactor modeling.

Since long lasting computational time is needed in order to achieve the steady state for this system (see Chapter 4 and Chapter 5) complications due to the use of complex model should be avoided; a global model is inserted in this study. According to [53] devolatilization is described by a simple competitive reaction model. The model also counts in the secondary reactions of tar by means of a first order reaction kinetics depending on tar concentration. A pseudo homogeneous description of the reacting system is assumed in this work.


50

3.3.2.2 Biomass

Though a heterogeneous description of the process should be addressed even in this case, the greatest part of works regarding biomass pyrolysis found in the literature deals with global models. One step or multiple step reaction mechanisms dealing with CRM or MRM and the exploitation of DAEM models are available [6, 11, 12, 13, 15, 26, 32, 39, 43, 46, 54].

Some works on pyrolysis describe the process kinetics as a function of hemicelluloses, cellulose and lignin content of the fuel [54].

Several models concerning with low temperature pyrolysis, at which the contribution of tar decomposition is negligible, do not account the secondary reactions of tar, most of them use a nitrogen flow during their experiments in order to remove and cool the gas as quickly as possible and split secondary from primary reactions.

Several correlations that allow us to calculate the gas production or the solid particle weight change are available, but some assumptions are necessary to use those models in order to calculate the gas and the char production from wood. Among the different global models present in the literature a pseudo homogeneous kinetic model based on competitive reactions is considered in this work [12, 32]. According to it we consider a two step process: a set of three first order reactions depending on the biomass content of the bed for primary pyrolysis and a first order reaction depending on tar concentration in the gas phase for secondary pyrolysis.

3.3.3 Combustion and gasification

The investigation of gas-char reactions should reckon with particle behavior, sorption and diffusion of reactants and products. An exhaustive description of the process is usually hindered by their interaction.

Nevertheless the influence of diffusion inside the pores is usually avoided during the study of a single particle behavior and the reaction rates obtained are thus in kinetic control. Three different kinetic models are proposed for char combustion: simple power law with reaction order up to one, Langmuir-Hinshelwood models and more complex three step global kinetic models [27]. The first is very simple but still finds its applications whenever a relation for restricted operating conditions range is needed. It assumes a one step reaction


51

from char to give CO or CO2. The second method assumes an intermediate step where an adsorbed oxygen

atom is supposed to exist. Within the second method a total conversion to CO is assumed.

In the third method products differentiation depending on the active site is assumed; the possibility to produce CO and CO2 exist in this case.

It is widely recognized that product distribution of carbon combustion between CO and CO2 is a function of temperature. It means that the third model is more appropriate. Whit the use of this model the transition from low to high order kinetics at different temperatures is as well possible and the kinetic expression has a greater validity. As previously pointed out all the kinetic models are developed under kinetic control. The results of these works is usually suitable for pulverized coal combustion models or fluidized gasification where small particles are used and kinetic control is more likely to occur. In the case of fixed bed gasification particle size ranges between 3-38 mm [7]. At these conditions diffusion resistance is for sure relevant and the previously mentioned kinetic expression could overestimate the reaction rate.

A shrinking core reaction model is adopted in this work and a first order reaction is considered. The reaction is considered to take place on the external surface of the unreacted core. The importance of the diffusion resistance inside the particle is, therefore, lower than the resistance due to the boundary layer, the kinetic and external diffusion resistance have to be considered during the reaction rate calculation.

The dependence from temperature of the CO/CO2 ratio is considered and calculated according to [2].

Coal char reactivity in a CO2 and H2O environment has been widely studied. Langmuir Hinshelwood type models seem to be appropriate to the description of this system. Nevertheless the most part of these considerations are related to a char type and cannot be applied to all of them. The most part of the kinetic models was indeed developed in particular conditions [19, 36, 40, 44, 55].

The presence of competition and inhibition effects between CO2 and H2O gasification is also suggested by some authors but the most part of works treat the two reactants separately [44].

The influence of thermal history of the char sample as well as its structure and the variation of char reactivity with conversion has been recognized in many works but no quantitative information on their effect has been given yet.


52

Some authors suggested that since some discordance still exists in the literature about the topic, no mathematical model is still capable to describe in detail the behavior of the gasification of a char particle [36]. A proper model should be chosen depending on the aim of the work. A shrinking core reaction model is adopted in this work; diffusion and kinetic resistance are considered [25].

The topic has been studied extensively even in the case of biomass char although less information is available in comparison with coal char. Even in this case the most part of authors report char reactivity for well defined operating conditions and thermal history of the char particle [44].

However, all the authors agree to say that biomass char is more reactive than coal char. The topic has received a wider attention especially in the last twenty years after that the suitability of a co-gasification process was reported [10, 57].

As reported in the literature char reactivity decreases from biomass to lignite and then decreases more in the case of a hard coal [29]. The different porosity, composition and different carbon structure are reported as the main reasons for this difference. The influence of the different ash content and its composition is also indicated as a possible cause [5]. An increased char reactivity was obtained from a biomass-coal blend with respect to coal char for a combustion system. The influence of biomass on hard coal char was lesser than that observed in case of a lignite- biomass blend. As reported in the literature a mixture of birch and Daw Mill coal showed almost twice as reactivity as the coal char alone, while this effect decreased if ash was removed.

A shrinking core reaction model is adopted in this work to describe biomass char reaction either for combustion either for gasification; diffusion and kinetic resistance are considered [25]. As previously discussed, it is difficult to find the proper kinetic parameters for the char we are working with and with the same thermal history; an ad hoc investigation should be carried out. For this reason the same model and the same kinetic parameters developed for coal char are here used. The different reactivity due to porosity differences is counted in while the effect of the different ash composition and different char structure is not considered by this model. In order to understand the influence of kinetic parameters and the importance of this assumption a proper sensitivity analysis is shown in Chapter 4.


53

Table 3: list of solid phase reactions and their respective kinetics law inserted in the model.

SOLID PHASE REACTIONS*

Ref. Drying

��(@'�AB → @'B6� # = ?D�E [9]

Primary pyrolisys

Wood

F��G → �� + �� + �� + ��H + �� # = ?I0J�I [47]

F��G → '6A< # = ?I0K�I [47]

F��G → /ℎ6A # = ?I0L�I [47]

Coal

��6� → �� + �ℎ6A # = � ?�MNJ

O

��=PQREST − PQR>�1 [53]

��6� → �� + �ℎ6A # = � ?�MNK

O

��=PQRK

EST − PQRK>�1 [53]

��6� → �� + �ℎ6A # = � ?�MNL

O

��=PUKREST − PUKR>�1 [53]

��6� → �� + �ℎ6A # = � ?�MNV

O

��=PUK

EST − PUK>�1 [53]

��6� → ��H + �ℎ6A # = � ?�MNW

O

��=PQUV

EST − PQUV>�1 [53]

��6� → '6AX + �ℎ6A # = � ?�MNY

O

��=PZS[\

EST − PZS[\>�1 [53]

Char gasification and combustion

6 � + �� → 2=^ − 1> �� + =2 − ^> �� # = 67�RK

12 ?X-�

+ 1?E

[25]

� + 2�� → ��H # = 67�UK

12 ?X-�

+ 1?E

[25]

� + �� → 2 �� # = 67�QRK

12 ?X-O

+ 1?E

[25]

� + �� → �� + �� # = 67�UKR

12 ?X-H

+ 1?E

[25]


54

Table 4: list of gas phase reactions and their respective kinetics law inserted in the model.

GAS PHASE REACTIONS

Secondary pyrolysys Wood

'6A< → �� + �� + �� + ��H + �� # = ?0W�ZS[_ [32]

'6A< → /ℎ6A # = ?0Y�ZS[_ [14]

Coal

'6AX → �ℎ6A + ��H + �� # = `X0a�ZS[\ [53]

Water-gas shift reaction

�� + �� → �� + �� # = `<-=�QR�UKR − �QRK�QUK?<-

> [16]

Gas phase combustion

2�� + �� → 2�� # = `-X��UKR.��RK.��QR [30]

A kinetic description of the reaction of the gas phase is as well counted in the model. In accordance with homogeneous rate expressions found in the literature all the secondary reactions of tar are inserted [30].

The methane combustion reaction was previously inserted. Nevertheless, as most

part of methane was produced once that oxygen had already been totally converted, its influence was found rather ineffective.

Carbon monoxide combustion is also inserted in the model; a complex homogeneous reaction rate is used to describe CO conversion. This reaction has been studied for a long time. There is a wide agreement between the authors with regard to the order of reaction and the competition effects in case of reaction happening in presence of steam. Nevertheless, authors disagree on the activation energy and pre-exponential factor and a wide range of values is present in the literature. The kinetic model used in this work is described in Table 4

An exhaustive description of all the reactions considered, with reference to the kinetic models used, is shown in Table 3 and Table 4.4


55

3.4 References

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2. Arthur J.R., Reactions between carbon and oxygen, Transactions of the Faraday Society, 47, 1951, 164-178.

3. Babafemi A. A., Hoanh N. P., Mathematical modelling of devolatilization of

large coal particles in a convective environment, Fuel, 74, 1995, 896-902. 4. Bird R. B, Stewart W. E., Lightfoot E. N., Transport phenomena, 1960,

New York, Wiley. 5. Brage C., Yu Q., Chen G., Sjöström K., Tar evolution profiles obtained from

gasification of biomass and coal Biomass and Bioenergy, 18, 2000, 87-91. 6. Branca C., Di Blasi C., Kinetics of the isothermal degradation of wood in the

temperature range 528–708 K, Journal of Analytical and Applied Pyrolysis, 67, 2003, 207–219.

7. Bridgwater A. V., The technical and economic feasibility of biomass gasification for power generation, Fuel, 74, 1995, 631-653.

8. Bryden, K.M., Kennet, W.R., Rutland, C.J., Modelling thermally thick

pyrolysis of wood. Biomass & Bioenergy, 22, 2002, 41-53. 9. Chan WR, Kelbon M, Krieger BB., Modeling and experimental verification of

physical and chemical processes during pyrolysis of a large biomass particle, Fuel, 64,

1985,1505–1513.

10. Collot A. G., Zhuo Y., Dugwell D. R., Kanndiyoti R., Co-pyrolysis and co-gasification of coal and biomass in bench-scale fixed-bed and fluidized bed reactors, Fuel, 78, 1999, 667-679.

11. Di Blasi C, Branca C, Santoro A, Perez Bermudez R. A., Weight loss

dynamics of wood chips under fast radiative heating, Journal of Analytical and Applied Pyrolysis, 57, 2001, 77–90.

12. Di Blasi C., Analysis of convection and secondary reaction effects within porous solid fuels undergoing pyrolysis, Combustion Science and Technology, 90, 1993,

315–40.

13. Di Blasi C., Comparison of semi-global mechanisms for primary pyrolysis of lignocellulosic fuels, Journal of Analytical and Applied Pyrolysis 47, 1998, 43–64.

14. Di Blasi C., Modeling and simulation of combustion processes of charring and non-


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charring fuels, Progress in Energy and Combustion Science, 19, 1993, 71–

104. 15. Di Blasi C., Physico-chemical processes occurring inside a degrading two dimensional

anisotropic porous medium, International Journal of Heat and Mass Transfer, 41, 1998,4139-4150.

16. Di Blasi, C., Dynamic behaviour of stratified downdraft gasifiers . Chemical Engineering Science, 55, 2000, 2931-2944.

17. Erincin D., Sinağ A., Misirlioğlu Z., Canel M., Characterization of burning and CO2 gasification of chars from mixtures of Zonguldak (Turkey) and Australian

bituminous coals, Energy Conversion and Management, 46, 2005, 2748-2761.

18. Froment G. F., Bischoff ,K. B.,. Chemical reactor analysis and design, 1979, New York, John Wiley & sons.

19. Fu W.B., Wang Q.H., A general relationship between the kinetic parameters for the gasification of coal chars with CO2 and coal type. Fuel Processing Technology, 72, 2001,63-77.

20. Gao, N., Li, A., Modeling and simulation of combined pyrolysis and reduction zone

for a dowdraft biomass gasifier. Energy conversion and management, 49, 2008, 3483.

21. Genetti D. An advanced model of coal devolatilization based on chemical structure, 1999, M.S. Thesis, BYU.

22. Gøbel, B, Henriksen,U., Jensen, T.K., Qvale,B., Houbak, N., The development of a computer model for a fixed bed gasifier and its use for optimization

and control. Bioresource Tecnology, 98, 2007, 2043-2052. 23. Gürüz G. A., Üçtepe Ü., Durusoy T., Mathematical modeling of thermal

decomposition of coal. Journal of Analytical and Applied Pyrolysis, 71, 2004, 537-551.

24. Hobbs M. L., Radulovic P. T., Smoot D., Combustion and gasification of coals in fixed-beds .Progress in energy combustion science, 19, 1993, 505-586.

25. Hobbs M. L., Radulovic P. T., Smoot D., Modelling fixed-bed coal gasifiers . AIChE Journal, 38, 1992, 681-702.

26. Hu S., Jess A., Xu M., Kinetic study of Chinese biomass slow pyrolysis: Comparison of different kinetic models, Fuel, 86, 2007, 2778-2788.

27. Hurt R. H., Calo J. M., Semi-global intrinsic kinetics for char combustion modeling, Combustion and Flame, 125, 2001, 1138-1149.


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28. Iniesta E., Sánchez F., García A.N., Marcilla A., Yields and CO2 reactivity of

chars from almond shells obtained by a two heating step carbonisation process. Effect

of two different chemical pre-treatments and ash content, Journal of Analytical and Applied Pyrolysis, 58-59, 2001, 983-994.

29. Kastanaki E., Vamvuka D., A comparative reactivity and kinetic study on the combustion of coal-biomass char blends, Fuel, 85, 2006, 1186-1193.

30. Kaushal, P,Proll, T., Hofbauer, H., Model development and validation: Co-

Combustion of residual char, gases and volatile fuels in the fast fluidized combustion

chamber of a dual fluidized bed biomass gasifier. Fuel,86, 2007, 2687-2695. 31. Kennet, W.R., Bryden, K.M., Ragland, W., Numerical modelling of a deep,

fixed bed combustor . Energy & Fuels, 10, 1996, 269-275. 32. Liden AG, Berruti F, Scott D.S., A kinetic model for the production of liquids

from the Bash pyrolysis of biomass. Chemical Engineering Communications,

65, 1988, 207–221.

33. Liliedahl T., Sjöström K., Modeling of coal pyrolysis kinetics, AIChE Journal, 40, 1994, 1515-1523.

34. Maniatis K., Beenackers A. A. C. M., Tar protocols. IEA Bioenergy gasification task, Biomass and Bioenergy, 18, 2004, 1-4.

35. Milne, T.A., Abatzoglou, N.,Evans, R.J., Biomass Gasifier “Tars”: their nature, formation and conversion. Report NREL/TP-570-25357, NREL, 1998, Golden, Colorado, USA.

36. Molina A., Mondragón F., Reactivity of coal gasification with steam and CO2 Fuel, 77, 1998, 1831-1839.

37. Moreea-Taha, Ruskana, Modelling and simulation for coal gasification . IEA Clean Coal Centre report, CCC/42 ,2000.

38. Niksa S., Predicting the devolatilization behavior of any coal from its ultimate analysis, Combustion and Flame, 100, 1995, 384-394.

39. ÂrfaÄo J.J.M., Antunes F.J.A., Figueiredo J.L., Pyrolysis kinetics of lignocellulosic materials: three independent reactions model, Fuel, 78, 1999, 349-358.

40. Ochoa J., Cassanello M. C., Bonelli P. R., Cukiermann A. L., CO2

gasification of Argentinean coal chars: a kinetic characterization, Fuel Processing Technology, 74, 2001, 161-176.

41. Paviet F., Bals O., Antonini G., The effects of diffusional resistance on wood char gasification, Journal of safety and environmental protection, 86, 2008,


58

131-140. 42. Qader S.A., Natural gas substitutes from coal and oil , 1985, Amsterdam,

Elsevier, 1985 43. Radmanesh R., Courbariaux Y., Chaouki J., Guy C. A unified lumped

approach in kinetic modeling of biomass pyrolysis, Fuel, 85, 2006, 1211–1220. 44. Roberts D.G., Harris D.J., Char gasification in mixtures of CO2 and H2O:

competition and inhibition, Fuel, 86, 2007, 2672-2678. 45. Salatino P., Senneca O., Masi S., Gasification of a coal char by oxygen and

carbon dioxide. Carbon, 36, 1998, 443-452. 46. Sonobe T, Worasuwannarak N, Kinetic analyses of biomass pyrolysis using the

distributed activation energy model, Fuel, 87, 2007, 414-421. 47. Thurner F, Mann U., Kinetic investigation of wood pyrolysis, Industrial and

Engineering Chemistry Process Design and Development, 20, 1981,

482–488.

48. Tinaut F. V., Melgar A., Pérez J. F., Horrillo A, (2008). Effect of biomass particle size and air superficial velocity on the gasification process in a downdraft

fixed bed gasifier. An experimental and modeling study. Fuel processing technology, 89, 2008, 1076-1089.

49. Ulloa ., Gordon A. L., García, Distribution of activation energy model applied to the rapid pyrolysis of coal blends, Journal of Analytic and Applied Pyrolysis, 71, 2004, 465–483

50. Wang, F. Y., Bhatia,S., A generalized dynamic model for char particle gasification with structure evolution and peripheral fragmentation. Chemical engineering science, 56, 2001, 3683-3697.

51. Williams A., Backreedy R., Habib R., Jones J. M., Pourkashanian M., Modelling coal combustion: the current position. Fuel, 81, 2002, 605-618.

52. Wu, Z., Fundamentals of pulverised coal combustion . IEA Clean Coal Centre report, CCC/95, 2005.

53. Wutti, R., Peter, J., Staudinger, D., (1996). Transport limitations in pyrolysing coal particles . Fuel, 75, 843

54. Yang H., Yan R., Chen H., Lee D.H., Zheng C, Characteristics of hemicellulose, cellulose and lignin pyrolysis, Fuel, 86, 2007, 1781–1788.

55. Ye D. P., Agnew J. B., Zhang D. K., Gasification of a south Australian low-rank coal with carbon dioxide and steam: kinetics and reactivity studies, Fuel, 77, 1998, 1209-1219.


59

56. Zhao Y., Hao W., Zhihong X., Conceptual design and simulation study of a co-gasification technology, Energy Conversion & Management, 47, 2006, 1416-1428.

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Part II

Model results

A MATHEMATICAL MODEL FOR A FIXED BED UPDRAFT REACTOR MODEL RESULTS

63

Introduction

The model has been tested in the case of biomass feedstock as well as coal feedstock; results obtained are discussed in this part of the thesis.

The system is intrinsically complex and the greatest part of time has been aimed to the numerical solution of governing equations.

Since a lack in the description of the dynamic behavior of the gasifier and a small sensitivity analysis is present in previous works, especially the goal to enlarge the range of convergence for some parameters was pursued.

A brief description of the numerical solution adopted in this work is presented in the following, while results obtained for biomass and coal feed are respectively shown in Chapter 4 and Chapter 5.

Numerical solution

The model consists in twenty-four differential and four algebraic equations. The system of governing differential equations is solved with the use of an explicit finite difference method.

The system has been under investigation for several years at our department, and many solution methods were used in the past to solve the equations.

A first solution of the steady state system was first attempted. The problem, due to the countercurrent flow of the two phases, was a Boundary Value Problem (BVP).

Conditions for the solution of the system are respectively given at the top of the reactor for the solid phase and at the bottom for the gas phase. Due to the BVP it is very difficult to find an optimal solution method and reach a quick calculation convergence. In [3] the system was solved with the LSODE package. The iterative method was an upward and downward integration. The first step was the integration of all the differential equations from the top to the bottom of the gasifier. The second was an upward integration of the gas phase equations from the bottom to the top of the gasifier at constant solid phase variables. Results of the upward integration step were prediction values for the second downward integration. Convergence criterion was the equivalence of the gas phase properties (composition and temperature) on the top of the gasifier for two consecutive iterations.


64

A steady state solution was previously obtained through the downward and upward integration of the mass and the heat balances by means of a Runge Kutta method.

In order to face all the problems related to the countercurrent system this method was coupled with the secant method and the simple substitution method. The next attempt was in that case calculated through the overall material balance for the whole reactor. At that stage a set of equations for the steady state was used to describe the system. Unluckily the solution was achieved for a restricted range of operating conditions. Main problems were connected to the complexity of the system, the presence of different order of magnitude between the kinetic rates of the processes involved and instability of the system.

In order to improve the stability of the system and to increase the versatility of the model a different approach is used at present. The governing equations for the gasifier are now the mass and the heat microscopic balances where the storage term is also considered. An investigation of the stationary behavior of the reactor is then possible by the use of the steady state achievement of the dynamic model.

Now we are still interested in the steady state conditions of the reactor but the approach to the problem has changed. Equations consider the unsteady state and the system has moved from a BVP of a ODE to a Initial Value Problem (IVP) of a Partial Differential Equation (PDE).

The gasifier is now divided in a set of cylindrical cells; each one has a radius equal to the reactor radius and a length equal to the integration step. The central finite difference method is used to build the grid while the Eulero method is used for time integration.

The system still has stability problems, but they can be more easily controlled in this way.

Differential equations were studied and a fictitious diffusivity of the gas phase was introduced to increase the system stability. The introduction of this parameter was necessary because the solid residence time is very high and a small integration step (high calculation time) would be required to reach the steady state condition without this tool.

In order to obtain the proper integration steps, the stability of parabolic differential equation is studied. Given the parabolic equation:


65

−bFTT + +FT + Fc = d − eF

For the linear case convergence of the numerical solution is related to the solution stability. Whenever the solution is stable, convergence is achieved and vice versa. In [1] are reported convergence criteria for the solution of a linear parabolic equation. In this study they are used in order to define the relation between the time integration step and the spatial integration step for a linear case associated to the non linear case under investigation. Integration steps obtained for the linear case were then employed as a reference for those concerning the non linear case. According to the convergence criteria reported in [1], when the numerical solution of the differential equation is obtained through an explicit method, the following relation exists between the two integration steps:

∆f ≤ 2b|+|

∆i ≤ ∆f�

2b + e∆f�

For a mass balance:

b = +,��

+ = �

e = 0 For heat balances:

b = ?�/0

+ = �

e =4G ℎ< + k�67

�/0

The value of the parameters A, D, and F were studied for the system under investigation. Kinetic rates were considered temperature and concentration independent and mean values were used to perform the analysis.

A short integration step owing to high calculation time was first obtained. The step used in the combustion section was especially found to influence stability.

A variable step along the reactor was the first solution to the problem,


66

nevertheless this method was ineffective and long lasting calculation times were still present.

The introduction of a fictitious diffusivity has been the key tool to decrease the calculation time. A spatial integration step of 0.05m and time integration step of 0.01s has been chosen at the end of the investigation in the case of biomass; 0.05 m and 0.001s are respectively used in the case of a coal feedstock.

Differently from other works [2, 4], high spatial integration steps are used here. Although this value is large in comparison with the reactor length, it was chosen in order to allow the study of the system for long times. The choice has to be seen as a compromise between the quality of results and the possibility to enlarge the range of convergence of the method.

In order to increase the stability of the system, material balances and energy balances are separated during calculation: a fixed temperature profile is kept while a number of integrations are carried out for mass continuity equations. Subsequently, a number of integrations with the fixed gas and solid phase composition previously obtained are carried out for energy balances. As soon as convergence for energy equations is reached a new integration cycle is carried out for material balances. The whole cycle is repeated until the whole system has reached the steady state condition.

Initial conditions

Initial conditions assumed in this work are in accordance with the start up of an existing coal gasification pilot plant. In that case the reactor is filled up with charcoal, a nitrogen flow is fed to the reactor and heat is provided to the gasifier with the aid of electrical heaters. As soon as the desired temperature profile is established, coal, air and steam are fed to the reactor. Nevertheless many decisions are left to the operator and every start up is different, the moment when the heaters are switched off depends on the run and the gas feed flowrate is often varied during operation.

The possibility to reproduce exactly the real start up conditions is far from being possible, an approximation of them is attempted in this work.

At the initial state the reactor is assumed to be filled up with char. A linear temperature profile is assumed in the bed: the char temperature ranges between 820°C at the bottom of the gasifier and 25°C at the top. The gas phase is assumed to be composed of pure nitrogen and a linear temperature profile


67

(ranging between 850°C at the bottom and 400°C at the exit) is assumed for the gas phase. Temperature values were chosen according to the real plant.

When the simulation starts, a step variation is introduced to the gas and the solid flowrate and composition. A gas flow constituted by air and steam and a coal feed are respectively fed to the bottom and to the top of the reactor. The system dynamics is studied starting from this moment.

References

1. Barba D., Calcolo elettronico nell’ingegneria chimica, Roma ,1971, Edizioni Scientifiche Siderea.


3. Hobbs M. L., Radulovic P. T., Modeling fixed-bed coal gasifiers, AIChE

Journal, 38, 1992, 681-702. 4. Mandl C., Obernberger I., Biedermann F., Updraft Fixed-bed gasification

of softwood pellets: mathematical modeling and comparison with experimental


Chapter

4

Biomass gasification

The mathematical model described in Chapter 3 has been used to study the behavior of an atmospheric gasifier fed with birch wood. Dynamics starting from well defined initial conditions is described in this Chapter. An overview of the steady state behavior corresponding to the operating conditions of a reference case is given. The influence of the air to fuel ratio, the steam to fuel ratio, reactor geometry, some kinetic parameters on the gasification process and the importance of a good description of the heat exchange inside the bed is discussed at the end of the Chapter.

CHAPTER 4 BIOMASS GASIFICATION

71

4.1 Introduction

Results concerning the numerical solution of the equations presented in Chapter 3 for a biomass feedstock are discussed in this Chapter; results are divided in three parts:

� the description of dynamic behavior of the gasifier starting from the initial conditions discussed in paragraph Initial conditions;

� the description of the steady state of the reactor; � a sensitivity analysis.

A reference case concerning the operating parameters given in Table 5 is used to discuss the dynamic and the steady state behavior of the gasifier.

If necessary for comparisons or a better explanation of the reactor behavior this case study is in the whole Chapter referred as the main case or the reference case and every change with respect to these values is pointed out.

Table 5: parameters used to run the simulation.

Parameter value unit

Inside diameter 0.3 m Bed height 0.5 m

Bed pressure 1.0 Atm Biomass feed 0.005 Kg/s

Particles diameter 1.0 cm Void fraction 0.5 - Ash content 0.5 % weight

Moisture content 8 % weight Solid feed temperature 25 °C Air feed temperature 287 °C

Steam feed temperature 287 °C Wall temperature 100 °C

Air-solid feed ratio (ATF) 0.9 Weight Steam-solid feed ratio (STF) 0.0 Weight

Table 6: elemental analysis of the wood used to run the simulation.

Elemental analysis wt%

C 49.0

H 6.2

N 0.3

O 44.5


72

Birch is the feedstock used, ultimate analysis of the fuel is given in Table 6.

4.2 Dynamic behavior

A short-time behavior, a long-time behavior and an ash layer formation above the grate are observed in the reactor.

Figure 1 contains the time variation of the temperature profiles of the gas and the solid phase for times up to 4000s.

During the short time the gas temperature decreases until it becomes lower than solid temperature in the whole reactor but in the biomass feeding section. The formation of the peak is occurring at the same time in the gas combustion section. The heat provided by the exothermic reactions is storing in that section: the heat produced is too low or it is not dragged to the sections above and it’s stored there. The gas and the solid temperature decreases so that the whole reactor is cooling off but in the combustion zone where a temperature peak is forming.

Because of the particular initial condition, reactor geometry chosen and because of the differences between the start up of a real plant and our

Figure 1: variation with time[s] of the solid temperature profile (solid line) and the gas temperature profile (dashed line) for the reference case.

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1.200

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]

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Initial 1000s 2000s 3000s 4000s


73

approximation, the difference between the peak temperature, previously reported in previous works as well, and the rest of the reactor where

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s fr

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Initial 1000s 2000s 3000s 4000s

Figure 3: time variation of the wood content profile (solid line) and the moisture content profile (dashed line) for the reference case.

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0,00 0,10 0,20 0,30 0,40 0,50

Mas

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Initial 1000s 2000s 3000s 4000s

Figure 2: time variation of the char content profile (solid line) and the ash content profile (dashed line) for the reference case.


74

temperature is very low is higher. As a consequence of the high spatial velocity (especially in comparison with

the solid phase) the dynamics of the gas phase is fast: after some seconds the gas composition reaches a level and does not experience relevant changes for short time (within minutes). As a result of the adaption to the new temperature profile and bed composition, further small variations are noticeable for longer times.

Since temperature decreases, drying and pyrolysis are slow up and the bed composition along the reactor changes. As shown in Figure 3, drying and pyrolysis sections lengthen and the char bed length decreases at the same time. Although in the gasification and combustion region the solid density is constant and the bed composition experiences just some small changes (see Figure 2 and Figure 3), the solid velocity profiles given in Figure 4 show that solid velocity decreases and the bed slows up due to the solid reaction: conversion has to be calculated on a flowrate basis in this system.

During the peak formation the solid temperature in the pyrolysis section decreases driven by the decreasing gas temperature and the difference between the two decreases.

The higher is the peak the higher is the driving force; the heat dragged to the

Figure 4: time variation of the solid velocity profile along the bed for the reference case.

0,0E+00

2,0E-05

4,0E-05

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1,0E-04

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/s]


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75

upper part of the reactor increases. Once that the peak reaches a maximum the solid temperature stops

decreasing in the rest of the gasifier and, except for the combustion section, the gas temperature becomes higher than the solid temperature.

Starting from this moment the peak decreases and heat is transferred to the upper part of the reactor. Therefore, in the endothermic reactions zone temperature of both, the gas and the solid phase, increases.

Some oscillations in the solid phase temperature and in the gas phase temperature are found during the long-time observation of the system. The major oscillations are in the temperature profile (see Figure 5 ).

In [2] is indicated the presence of some oscillations in the combustion zone of a biomass fed updraft reactor. Authors in that case obtained periodical oscillations and gave a physical explanation of the phenomenon. Nevertheless, they could not prove with experimental evidence the existence of this behavior in a real system. They also indicated that they could not exclude with certainty that oscillations were due to the algorithm used for calculations.

In their work, the moment that oscillations started and the maximum time achieved in their simulations is also absent. Comparison between the two works is, therefore, difficult.

Figure 5: long time dynamics of the combustion peak for the reference case.


76

In the case presented in this Chapter oscillations do not show any periodicity. Some oscillations in the solid phase are observed for long time (with a recursion

Figure 7: long time dynamics of the combustion peak for the reference case

Figure 6: long time dynamics of the combustion peak for the reference case.


77

between two peaks in the order of twenty minutes). Almost the same oscillations are observed in the gas temperature (see Figure 6). As discussed in the paragraph regarding the numerical solution, the system solution is affected by instability problems. Therefore, oscillations are related to the process description and the algorithm chosen for calculation. Nevertheless, the rest of the reactor does not seem to be affected dramatically by this phenomenon and just small fluctuations are observed in the bed as well as in the gas composition and their importance is negligible.

Since the solid speed inside the reactor is very low especially at the bottom, where velocity is decreasing due to conversion and the solid is almost stationary, a long time (in order of hours) is required to obtain a change in composition or height of the reacting bed.

As shown in Figure 8, once that char conversion is approaching the complete consumption the solid velocity decreases rapidly and the bed composition changes in the combustion section. The birth of an ash layer is observed in this zone (see Figure 9).

This phenomenon is clearer when a value of the air to fuel ratio higher than the typical gasification range is used to run the simulation (ATF > 2.7). In

Figure 8: dynamics of the solid velocity profile in the combustion region for the reference case.

0,0E+00

5,0E-06

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1,5E-05

2,0E-05

0,40 0,42 0,44 0,46 0,48 0,50

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/s]


20000s 30000s 40000s 50000s


78

particular, for an air to fuel ratio equal to 4 the ash layer increases with time and, as shown in Figure 10, the front of combustion, going up, moves along the reactor. A different dynamics is also observed in this case. Since a large amount of oxygen is fed to the reactor the combustion zone is wider and higher temperature exists in the whole reactor. In this case the pyrolysis front does not move downwards as more as in the reference case. Moreover, after a small movement to lower sections it goes up again and at steady state the reacting bed becomes very short, the combustion section is limited and char reacts almost instantaneously once produced from pyrolysis.

Since a bed cooling off for a particular combination of operating parameters is reported in [1], the system has been tested for different conditions.

A range between 0.8-4 of the air to fuel ratio (ATF) and 0-0.6 of the steam to fuel ratio has been investigated.

For an air to fuel ratio lower than 0.8 the system is not able to ignite and the steady state obtained is an extinguished reactor. The gasifier is capable to ignite for all the values of the ATF ratio investigated up to 4. As reported in [2] this is in accordance with experimental evidence.

Moreover, the investigation shows that extinction of the bed is possible with

Figure 9: ash layer formation in the combustion region for the reference case.

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79

a proper amount of steam; as reported in Table 7 the higher is the ATF ratio the more steam is needed to cool off the reactor.

Since no extinction of the bed is observed for high values of the ATF ratio

Figure 11: time variation of the char profile for ATF=4.

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- 0,10 0,20 0,30 0,40 0,50

Mas

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Initial 4000s 8000s 14000s 18000s

Figure 10: ash layer formation in the combustion region for ATF=4.

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Initial 4000s 8000s 14000s 18000s


80

(high gas phase velocities) the effect of oxygen dilution due to steam is the most probable reason of the phenomenon.

Table 7: variation of STF at extinction with ATF

ATF STF at extinction

0.8 0.35 0.9 0.45

1.0 0.65

1.1 0.65

4.3 Steady state description

The steady state temperature profiles for the reference case are reported in Figure 12. Temperature has a monotone trend in the greatest part of the reactor but the presence of a peak very close to the grate, in the combustion section. The heat dragged by the gas phase from the combustion zone to the upper part of the gasifier is not enough to ensure a high temperature environment for a long fraction of the bed.

The gas heating section is very short; the ash temperature at the exit is

0

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400

600

800

1000

0,00 0,10 0,20 0,30 0,40 0,50

Tem

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]


Solid phase Gas phase

Figure 12: steady state temperature profile along the bed for the reference case


81

293°C. As expected, the gas phase temperature is driving the system in the upper

part of the bed and the biomass heating section appears larger than expected (10 cm). Temperature of the gas product is 144°C.

As temperature is lower than 300°C for more than a half of the reactor the most part of reactions is expected to take place in the last part of the gasifier.

The variation of the solid composition along the bed is shown in Figure 13. The bed is constituted by a solid heating section (0.1 m), a drying and pyrolisys section (for a total length of 0.30 m) and a gasification and combustion zone (0.10 m). Ash content of the bed is higher in the last part of the reactor where char is consumed.

As the bed composition is constant in the first part of the gasifier a large fuel heating section (10 cm) exists.

Drying is very slow, more than expected; drying and pyrolysis are sequential. Moreover, since the total wood conversion is obtained in 10 cm pyrolysis seems to be faster in comparison with drying. As temperature is driving the system the slowness of the phenomenon is mainly due to temperature limitations.

The combustion and the gasification processes are overlapped and char reacts very fast.

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Mas

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Wood Moisture Char Ash

Figure 13: steady state composition of the bed for the reference case.


82

Gasification and combustion seem to be faster than expected; two can be the main reasons: the unsuitability of the kinetic parameters used or a low solid rate fed to the reactor. However, a sensitivity analysis, presented in paragraph 4.4.3 showed just a small influence of the gas composition of the kinetic parameters.

As previous works indicated the presence of a short biomass heating, drying and pyrolysis section and a long char heating section results presented in this paragraph partly disagree with them. Nevertheless, they use different reactor dimensions and some differences in the reactor handling during the transient behavior could have influenced the steady state achievement.

According to the bed composition char conversion seems to be very low but, although the char content of the bed in the last section is 0.82, the solid conversion is very close to 100%. As shown in Figure 14 the solid flowrate is indeed decreasing with the reactor height and approaches almost zero at the bottom.

To comment the gas behavior a partitioning of the reactor in three regions is necessary: the combustion and gasification zone, the biomass pyrolysis and drying zone and the gas heating zone.

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Sol

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s]


Figure 14: solid flowrate profile along the reactor for the reference case


83

In accordance with previous works the combustion section is very short: as shown in Figure 15 the most part of carbon monoxide and carbon dioxide is

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tar CO CO2 H2 CH4 H2O

Figure 16: gas phase composition profiles in the endothermic reactions region.

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Mol

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CO CO2 H2 CH4 O2

Figure 15 gas phase composition profiles in the combustion region.


84

formed in this section; methane and hydrogen formation is very low. Because of pyrolysis reactions the gas phase is enriched in methane and

Figure 18: product gas composition for the reference case.

Figure 17: gas phase composition profiles in the biomass heating section.

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CO CO2 H2 H2O


85

hydrogen. Steam is also released in this section and the gas composition is affected by the water gas shift reaction. The carbon monoxide level decreases and the carbon dioxide level increases. This phenomenon is more evident during the slow drying process when steam released affects composition.

In the gas heating section the water-gas shift is the only reaction taking place in the gasifier.

The carbon monoxide content in the gas phase decreases along this section; the carbon dioxide and the hydrogen content increases instead. In accordance with the reactor type, results confirm that a large amount of tar (9.8%) is present in the product gas and a large amount of nitrogen (40%) and steam (13%) strongly dilute the gas. The gas composition obtained at the gas exit is given in Figure 18.

4.4 Sensitivity analysis

For a deeper understanding of the gasification process a sensitivity analysis of the model is presented in the following. The investigation deals with: the air to fuel ratio, the steam to fuel ratio, the reactor dimensions influence the heat exchange between the two phases description and the pre-exponential factor of the gasification reactor.

4.4.1 Influence of the Air To Fuel ratio (ATF)

The value of ATF is varied between the minimum allowing bed ignition and 4. When the air to fuel ratio ranges between 0.8 and 1, strict correspondence with biomass conversion is observed, the more air is fed to the reactor the higher conversion is achieved (see Table 8). Char combustion reactions are faster than gasification reactions and oxygen always reaches the complete conversion. It is summed up that the most part of conversion in the reactor is due to the char and the gas phase combustion while gasification reactions play a secondary role.

When ATF is higher than typical gasification value (the higher bound is about 2.26 for this fuel) an influence on the bed length is noticeable. The transient behavior shows wider combustion regions when the ATF is so high. As a consequence, the char bed present at the initial condition is converted and an ash layer that slowly increases moving the combustion front upward is


86

formed at the bottom of the reactor. As shown in Figure 19 at steady state the oxygen surplus is leaving the reactor unconverted while the combustion region

0,06

0,08

0,10

0,12

0,14

0,16

- 0,10 0,20 0,30 0,40 0,50

Oxy

gen

flow

rate

[Kg/

s]


Figure 20: solid and gas temperature profiles for an air to fuel ratio equal to 4.

0

180

360

540

720

900

0,00 0,10 0,20 0,30 0,40 0,50

Tem

pera

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[°C

]



Figure 19: oxygen flowrate variation along the bed for an air to fuel ratio equal to 4.


87

Table 8: solid conversion dependence on the air to fuel ratio.

ATF Conversion

0.8 0.944

0.9 0.958

1 0.971

and the whole length is very short. Figure 20 shows that, in this case, the peak moves to the upper part of the

reactor and the long heating section is absent. There is valuable evidence that temperature has a key role in this system and that processes mostly depend on temperature profiles established along the reactor.

The integration step is in this case too large to describe the system and the last observations are pointed out for their qualitative more than quantitative information.

4.4.2 Influence of the Steam To Fuel ratio (STF)

The investigation considers a variation of the steam to fuel ratio between zero and the value that caused the reactor extinction (in the following referred as maximum value). The most important variation observed in this case is in the gas composition.

The steam fed to the reactor affects the system in two different ways: it contributes to reaction and it dilutes the gas.

The first effect is more appreciable at low STF values. As shown in Figure 21, because of the gas-shift reaction, the H2 content of the gas phase and the CO2 level increase, the CO concentration decreases at the same time.

For higher STF values the dilution effect prevails: the H2O content of the gas increases at expenses of the other gases. The H2 level as well as CO2 level and CO content of the gas phase decreases. The double effect of steam injection is accordance to some experimental information (M. Siedlecki – private communication – Delft University) as a small amount of steam is able to increase the H2 content of the syngas while higher amounts are ineffective.

At high values of STF, the more steam is fed to the reactor, the more conversion drop is observed (see Figure 22). Since oxygen is diluted combustion reactions are slower and a higher char content is present in the solid phase at the


88

grate. As previously discussed, for a STF value higher than 0.45 the bed is not

capable to ignite. Since the sensitivity analysis shows that the dilution effect

Figure 22: steam to fuel ratio influence on solid conversion.

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CO CO2 H2 CH4 H2O tar

Figure 21: steam to fuel ratio influence on the gas composition.


89

plays an important role at high steam contents this is the most reliable reason of the phenomenon.

4.4.3 Influence of gasification reactions

The kinetic parameters inserted in the model for gasification reactions are given for coal char. As biomass char reactivity is recognized to be higher than coal char reactivity, the effect of an increase of the kinetic rate is considered in order to investigate the importance of this assumption.

Small and large variations are introduced to t he values of pre exponential factors of the kinetic constants in order to

reduce the kinetic resistance. Their value is multiplied by 2, 3 and 10. Results concern the contemporaneous increase of the three gasification rates, their synergic effect is shown.

The kinetic parameters do not affect dramatically the system: in agreement with the previous consideration the combustion reaction is the most important in this system.

Temperature profiles do not differ between the cases. Figure 23 contains the variation of the gas composition. An increase of the

0,000,020,040,060,080,100,120,140,160,180,20

CO CO2 H2 CH4 H2O

Mol

ar fr

actio

n in

the

gas

phas

e

Multiplicative coefficient =1 Multiplicative coefficient =2

Multiplicative coefficient =3 Multiplicative coefficient =10

Figure 23: influence of gasification reactions on the product gas composition.


90

CO content is noticeable; steam, as a reactant in the gasification reactions, decreases; CO2 and H2 levels, mainly produced in the water-gas shift reaction, remain almost unchanged; a little increase in the methane content is observed.

4.4.4 Influence of the gas-solid heat exchange

As the correct description of heat exchange is found to be an important issue on the gasifier modelling, the multiplicative coefficient ξ is varied in order to improve the gas-solid heat exchange. Moreover, the presence of reaction introduces some uncertainties on the heat transport description and a sensitivity analysis is needed in order to understand how much instrumental is this parameter.

The most important effect is observed in the temperature profile. As shown in Figure 24 the higher is the value of ξ the better heat exchange exists between the phases and lower differences exist between the gas and the solid temperature.

Since in the combustion region radiation between the phases plays an important role and the two temperatures are already very close, this is especially

Figure 24: gas-solid heat exchange term influence on the solid (solid line)and the gas (dashed line) temperature profile.

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0,00 0,10 0,20 0,30 0,40 0,50

Tem

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]


ξ= 0.8 ξ=0.6 ξ=0.40


91

true in the low temperature region. When ξ is equal to 0.8 the high temperature region enlarges and the value of

the combustion peak decreases. As a consequence of the different temperature profile, some fluctuations in

the gas phase are also obtained.

4.4.5 Influence of the initial bed height

The influence of the initial bed length on bed composition and temperature profiles along the reactor is respectively shown in Figure 25 and Figure 26. Two situations are in this case compared: the reference case and an initial bed length equal to 1m.

Profiles in the two cases are similar. They mainly differ in the biomass heating section. The variation of the solid phase composition along the bed is almost the same in the two situations but in the combustion region. As shown in Figure 26 temperature profiles are not affected but a small difference in the combustion peak height is noticeable.

Figure 25: bed composition profiles along the bed for initial bed height respectively equal to 1m (solid lines) and 0,5m (dashed lines)

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Wood Moisture Char Ash


92

Reactor performance is more related to the operating conditions than the initial reactor length.

4.5 Conclusions

The model described in Chapter 3 has been used to study the behavior of an updraft fixed bed gasifier where birch wood is used as feed.

The dynamic behavior and the steady state condition of a main case is described in this Chapter.

It is observed that the gas phase dynamics is faster than the solid phase dynamics: the first is in the order of minutes, the second is in the order of hours.

Short time effects are discussed during the study of the transient behavior, but the most important information is obtained from a long time analysis.

Bed extinction is observed when a high amount of steam is injected to the reactor. The particular amount is related to the air to fuel ratio used, the higher is ATF the higher steam amount is needed for bed extinction. Since reactor injection is possible for the whole range of air to fuel ratio investigated the

0

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960

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0,00 0,20 0,40 0,60 0,80 1,00

Tem

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0,5 m 1,0 m

Figure 26: initial bed height influence on the solid(solid line)and the gas (dashed line) temperature profile.


93

dilution effect is indicated as the main reason of the bed extinction. Some temperature oscillations related to the numerical method chosen are

observed in the combustion region. Nevertheless they do not affect critically the system and can be neglected during the study.

At steady state a temperature peak is present in the combustion region while temperature is very low in the rest of the reactor. For this reason the biomass heating section, the drying section and the pyrolysis section are longer than expected. Combustion and gasification take place in a short region and an ash layer exist above the grate.

During the sensitivity analysis the importance of the heat produced and dragged to the upper part of the reactor is pointed out. The air to fuel ratio, the steam to fuel ratio and the heat exchange coefficient are the main parameters affecting the reactor performance.

A correspondence between the air to fuel ratio and the solid phase conversion is observed. On the contrary, the steam to fuel ratio is found to affect the gas composition. In particular, the injection of small amounts of steam increases the hydrogen content of the syngas.

Gasification reactions show a little influence on the reactor performance but affect the gas composition.

The initial bed height is almost irrelevant on the reactor performance since the reaction zones length and profiles shape is more related to all the other operating parameters. Further work is needed for this model. Model validation and calibration should be at this point carried out. A different calculation method should also be used in order to avoid every stability problem.

4.6 References






94


5. Yoon H., Wei J., Denn M. M., Transient behavior of moving-bed coal gasification reactors, AIChE Journal, 25, 1979, 429-439.

Chapter

5

Coal gasification

The model described in Chapter 4 has been used to study the behavior of an atmospheric gasifier fed with Pittsburgh coal seam n°8. The differences observed between coal and biomass gasification are discussed in this section, especially with reference to the dynamic behavior of the system. An overview of the steady state behavior corresponding to the operating conditions of a reference case and the influence of some operating parameters are discussed

CHAPTER 5 COAL GASIFICATION

97

5.1 Introduction

A case study concerning a coal gasification system is presented in this Chapter. Results are divided in three parts:

� the description of the dynamic behavior of the gasifier starting from the initial conditions discussed in paragraph Initial conditions;

� the description of the steady state of the reactor; � sensitivity analysis.

In particular the first two parts respectively refer to the dynamic behavior and the steady state behavior of the gasifier for the operating conditions shown in Table 9.

If necessary for comparisons or a better explanation of the reactor behavior

Table 9: operating conditions for the reference case.

Parameter value unit

Inside diameter 0.3 m Bed height 0.5 m

Bed pressure 1.0 Atm Coal feed 0.01 Kg/s

Particles diameter 1.0 cm Void fraction 0.4 -

Moisture content 10 % weight Solid feed temperature 25 °C Air feed temperature 287 °C

Steam feed temperature 287 °C Wall temperature 100 °C

Air-solid feed ratio (ATF) 4.0 Weight Steam-solid feed ratio (STF) 0.2 Weight

Table 10: proximate and ultimate analysis for the fuel used to run the simulations

Proximate analysis wt% dry basis

Ash 13.70

Fixed carbon 51.90

Volatile matter 34.40

Elemental analysis wt% dry basis

C 71.90 H 4.07 O 7.00. S 1.36 N 1.36


98

this case study is in the whole Chapter referred as the main case or the reference case and every change with respect to these values is pointed out.

A Pittsburgh seam n°8 coal is used to run the simulations, proximate and ultimate analysis for this fuel is given in Table 10.

5.2 Dynamic behavior

In this paragraph is presented the dynamic evolution of the system starting from the initial conditions previously described.

Figure 27 contains the time variation of the temperature profile of the system under investigation.

At the beginning the gas temperature decreases and becomes lower than solid temperature in the whole reactor but the combustion zone where the formation of a combustion peak is observed. The same behavior is described in Chapter 4 for a biomass feedstock.

Differently from a real coal gasification start up, where electrical heaters maintain the reactor in the proper thermal conditions until the complete bed ignition, during this simulation it is assumed that once that coal is fed to the

Figure 27: variation with time[s] of the solid temperature profile (solid line) and the gas temperature profile (dashed line) for the reference case.

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reactor heaters are switched off and the bed is capable to sustain reactions by itself. From the calculations it is observed that heat produced during exothermic reactions is accumulating in the combustion section and that the gas phase is not able to drag heat to the upper sections.

The gas phase is faster than the solid phase in the steady state achievement. It is indeed observed that the gas phase composition changes very fast: after some minutes it does not change substantially and further variations are related to its adaption to the new bed composition and temperature profile.

As temperature decreases in the pyrolysis and drying section, the reaction rate is slowing down and the pyrolysis reaction front moves downwards. The two zone are lengthen and a coal heating section is formed in the upper part of the gasifier.

Char conversion is initially very low but increases with time. Differently from the biomass case, solid composition and spatial velocity decrease at the same time in this system. The difference is due to the different solid density and to the different char amount produced from the two systems. It is indeed well known that fixed carbon is higher in the case of coal and a larger amount of char is released during its thermal decomposition.

Since both, solid velocity and composition, are changing in the reactor conversion for this system has to be calculated in a flowrate basis.

Temperature in the reactor decreases with time until the gas phase temperature becomes higher that the solid phase temperature. In fact the combustion peak is becoming higher while temperature is decreasing in the rest of the reactor: the higher is the peak the higher is the driving force and the heat dragged to the upper part of the reactor. Once that the peak reaches a maximum, the temperature depletion stops in the rest of the gasifier and the gas temperature becomes higher than the solid temperature except for the combustion section.

As previously discussed, oscillations of the temperature peak are observed for a biomass feedstock, instability of the system is the reason suggested in Chapter 4.

Different is the case of a coal feedstock. As shown in Figure 28 temperature in the combustion zone increases with time, heat produced in the combustion section is stored in this zone and it is not dragged to the upper part. The combustion peak is forming during that time and after 5 hours the highest temperature value is achieved. From this moment it is possible to observe just


100

small variations of the peak height and oscillations are totally absent. The different behavior between the two systems is related to the fact that a

smaller integration step is used in this case and the energy balance is more stable.

For a particular set of operating conditions a bed cooling off was observed in [1]. In order to verify the presence of this tendency, the air to fuel ratio (ATF) and the steam to fuel ratio (STF) are varied in a wide range.

An air to fuel ratio equal to 1 is the lower value adopted for the analysis while 6 is the maximum (according to [4] in a typical countercurrent coal gasification system ATF is in the range: 2.4-7.15). The investigation is carried out for different values of STF ranging from 0 to the highest possible for every value of ATF

For an air to fuel ratio lower than 1.3 the system is not able to ignite and the steady state obtained is an extinguished reactor. The gasifier is capable to ignite for all the values of the ATF ratio investigated up to 6. As reported in [2] this is in accordance with the experimental evidence.

Since no extinction of the bed was observed for high values of the ATF ratio, and therefore high gas phase velocities, the effect of oxygen dilution due to steam is the most probable reason of the bed cooling off.

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Figure 28: combustion peak dynamics for the reference case.


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5.3 Steady state description

The description of the steady state behavior of the gasifier for the main case is here presented.

Figure 30 shows the gas phase and the solid phase temperature profiles along the reactor.

As in the case of a biomass feedstock, the typical combustion peak exists in the last part of the gasifier. Moreover, since 1370 °C are achieved the height of the peak is greater in this case. In the rest of the reactor temperature is lower than 600 °C and has a monotone trend.

The gas heating section seems to be very short and the ash temperature at the exit is 340°C.

Temperature of the gas product is 103°C, lower than expected for this system. The reason can be found on the fact that the heat dragged by the gas phase from the combustion zone to the upper part of the gasifier is not enough to ensure high temperature for a long fraction of the bed.

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Figure 29: time variation of the coal content profile (dashed line) the moisture content profile (solid line) and char content profile(dots) for the reference case.


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Because of the low temperature region existing in the gasifier, pyrolysis is expected to take place slowly.

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Figure 31: steady state composition of the bed for the reference case.

Figure 30: steady state temperature profile along the bed for the reference case

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The bed is constituted by a solid heating section (0.1 m), a drying and pyrolysis section (for a total length of 0.30 m), a gasification and combustion zone (0.10 m) and an ash in the last part of the reactor where char is consumed.

As shown in Figure 31 a large fuel heating section (10 cm) is present in the be d.

Drying and pyrolysis are overlapped and very slow. Temperature is limiting reaction and in comparison with biomass pyrolysis coal devolatilization is slower.

Combustion and gasification are overlapped and char reacts very fast. As previously discussed in the case of a biomass feedstock (see Chapter 4) a

short solid heating, pyrolysis and drying section were found in previous works, nevertheless, the different reactor dimension is a possible reason for this difference.

Since a solid phase with high char content is leaving the reactor, conversion seems to be very low. This is actually false since conversion in this system has to be calculated in a flowrate basis. A conversion equal to 77 % is obtained in the reactor.

Gasification and combustion seem to be faster than expected. Since the most part of char conversion is obtained while oxygen is present in the gas phase (see

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CO CO2 H2 H2O O2

Figure 32: gas phase composition profiles in the combustion region.


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Figure 31 and Figure 32), combustion reaction seems to be more relevant than gasification reactions in this system. The reason of this short gasification and

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CO CO2 H2 H2O

Figure 34: gas phase composition profiles in the fuel heating section.

Figure 33: gas phase composition profiles in the endothermic reactions region.

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combustion zone is related to the low temperature existing in the rest of the reactor. From the sensitivity analysis is indeed obtained that the more heat is dragged to the upper part of the reactor the wider becomes this zone. Figure 32, Figure 33 and Figure 34 respectively contain the gas behavior in the gas heating section, the combustion and gasification zone, the pyrolysis and drying zone and the gas heating zone.

In accordance with previous works, where short combustion sections were found, the length of this section is 0.10m. A large amount of carbon monoxide is formed in this section: at this temperature CO is favored. Carbon dioxide is formed in this section but in a small amount, hydrogen formation is very low as well. Methane is almost absent, it is supposed to be produced from gasification reactions but, as previously discussed, the combustion reaction is dominating and the hydrogen gasification reaction, which is supposed to be the slowest, is the most affected from this phenomenon.

As expected coal devolatilization yields a large fraction of solid char and just small amounts of gas. Gas composition is subjected to small changes in the pyrolysis section. The water gas shift reaction is also taking place since the carbon monoxide level and the carbon dioxide level decrease as the hydrogen content increases.

In the last 0.10 m of its way along the reactor the gas composition is almost constant for the whole section that appears like a simple coal heating region.

Differently from the case of biomass tar content is zero in this case. The reason has to be found on the kinetic rate, too low in comparison with the CO formation reaction. In accordance with the reactor type, a large amount of nitrogen (50%) and steam (15%) strongly dilute the gas. The gas composition obtained at the gas exit is given in Figure 35.

5.4 Sensitivity analysis

A sensitivity analysis concerning the air to fuel ratio (ATF), the steam to fuel ratio (STF), the bed height and the solid flowrate is carried out and discussed in order to find the key parameters affecting the gasification process.

The influence of the heat transport along the bed is indicated as an important phenomenon for this system, nevertheless, its influence is considered the same for both, biomass feedstock as well as coal feedstock since the process is the same, it is thus assumed that its effect is mainly related to the system and is not


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investigated in this Chapter. The analysis of this parameter is already discussed in paragraph 4.4.4.

5.4.1 Influence of the Air To Fuel ratio (ATF)

The air to fuel ratio is varied from zero until 6. In the dynamic description of the system presented in paragraph 5.2 it is pointed out that the system does not ignite for low values of ATF. The range analyzed in this paragraph is 3-6.

In the description of the steady state behavior (see paragraph 5.3) the importance of char combustion more than char gasification reactions for this system is pointed out. The high temperature zone is restricted to the combustion region and the whole char conversion is achieved in that section. Is then clear why the air to fuel ratio plays such an important role. Since more oxygen is fed to this region, more char reacts by means of combustion reactions and the higher conversion is achieved (Figure 36).

ATF also affects the temperature profile. The more char is burnt the higher heat amount is released in this region: a small fraction is dragged to the upper part while the rest stores in the surroundings. A higher peak is obtained with an increase of this parameter (see Figure 38)

Figure 35: product gas composition for the reference case


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An influence of the composition in the combustion region is also observed. The higher conversion reflects on a higher ash content of the bed at the bottom.

Figure 37: bed composition dependence on the air to fuel ratio in the combustion region

Figure 36: solid conversion dependence on the air to fuel ratio.

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Char Ash


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Increasing the ATF it increases as well. When ATF is equal to 6 the ash layer is completely formed and the reaction front comes up to the sections above.

5.4.2 Influence of the Steam To Fuel ratio (STF)

Differently from the case presented in Chapter 4, in the case of coal feedstock, steam is always injected to the reactor in order to allow a better temperature regulation.

Two are the main effects connected to steam injection: the dilution effect and participation to reactions.

Steam is a reactant for gasification reactions as well as gas phase reactions (e.g. water-gas shift).

According to Qader [4] a large amount of steam is usually injected to this system. As reported is paragraph 5.2 a maximum value of STF, connected to the possibility to ignite the bed, exist and our sensitivity analysis therefore starts with small amounts but the parameter is tested just until that maximum.

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Figure 38: solid temperature profile (solid line) and gas temperature profile (dashed line)dependence on the air to fuel ratio


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Correspondence between STF and the syngas composition is observed. The more steam is fed to the reactor; the higher hydrogen is yielded by the process.

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Figure 40: steam to fuel ratio influence on solid conversion.

Figure 39: steam to fuel ratio influence on the gas composition.


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In Figure 39 is presented the variation of the gas composition with the parameter.

The CO level of the gas decreases with STF while the increase of hydrogen amount is observed at the same time. Nevertheless, the difference in the hydrogen content is high between the absence of steam and the injection of a small amount while it increases slowly when this parameter is increased. The increase is due to the water-gas shift reaction. Since the variation of syngas temperature is very small, even the syngas composition at equilibrium feels just a small effect. The dilution effect becomes relevant at high STF values. When the steam level is very high the concentration of the other gases indeed decreases. As previously discussed (see paragraph5.2), for a steam to fuel ratio higher than 1.2 the dilution effect prevails and the bed is not capable to ignite anymore.

Conversion and the height of the peak are also affected by STF. The higher amount of steam is fed to the reactor the lower is the peak. As temperature decreases in this section, it is expected to widen. The greatest difference is observed between the absence of steam and steam injection. Since a wider combustion region is formed even conversion increases. When the steam to fuel ratio is increased this effect still takes place but there is a competition with the temperature decrease in the region and the latter is more relevant: conversion depletion with steam content of the gas feed is observed (see Figure 40).

5.4.3 Influence of the solid flowrate

The effect of the solid flowrate is presented in this paragraph. As shown in Figure 42 the temperature profile does not differ substantially between the cases. A wider combustion region exist in the reactor for lower values of this parameter.

Although conversion increases with coal flowrate, bed ignition is difficult from a certain point and bed extinction is observed when a flowrate higher than 0.005 Kg/s is fed to the reactor.

5.4.4 Influence of the reactor length

The initial bed height is varied in order to understand its effect. Three tests are carried out: 0.5m, 0.75m and 1m.

In the main case pyrolysis and gasification are sequential; when the initial


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height of the bed is wider a char heating section exists in the gasifier. The higher is the initial bed height the longer is the heating section at steady

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Figure 42: solid conversion variation with coal flowrate.

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Figure 41: solid temperature profile (solid line) and gas temperature profile (dashed line) dependence on the solid flowrate.


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state condition; pyrolysis occurs at lower temperature when the initial bed is higher but devolatilization zone has almost the same length (see Figure 44)

Solid conversion and the width of the gasification and combustion region do

Figure 44: initial bed height influence on of the coal content profile (dashed line), the moisture content profile (solid line) and the char content profile (dots).

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Figure 43: initial bed height influence on the solid phase temperature profile (solid line) and the gas phase temperature profile (dashed line)


113

not seem to be affected by a different bed length and the profiles are overlapped in the three cases.

Gas composition does not differ between the cases. When the initial bed moves from 0.5m to 0.75m, the peak is less prominent, temperature profiles exhibit some small changes and a wider high temperature region is present is the gasifier.

5.5 Conclusions

Results concerning the steady state behavior, dynamics and a sensitivity analysis of a fixed bed gasifier, based on the numerical solution of the mathematical model previously described, are discussed in this chapter.

Since a difference in the spatial velocity of the two phases along the gasifier exists, the dynamic of the gas phase is faster than the solid phase dynamics, every change in the solid is gathered from the gas phase which quickly adapts to the new conditions.

The most important result concerning the transient behavior of the reactor is related to a long term analysis.

The system is able to ignite for a number of combinations of the air to fuel ratio and the steam to fuel ratio.

Typical operating conditions for the system are investigated. In particular, a bed extinction is observed when a high amount of steam is injected to the reactor. Since the sensitivity analysis showed prevalence of the dilution effect for high values of the steam to fuel ratio, the same is indicated as the possible reason of bed extinction.

Differently from the case of a biomass feedstock, presented in Chapter 4, temperature oscillations are absent for a coal feedstock.

At steady state a temperature peak is present in the combustion region while temperature is very low in the rest of the reactor. Long drying and pyrolysis regions are observed and a fuel heating zone exists in the reactor. Combustion and gasification take place in a short region and an ash layer is formed above the grate.

The importance of the heat produced and dragged to the upper part of the reactor in the reactor has to be underlined but its influence is neglected here since a sensitivity analysis is presented in Chapter 4. The air to fuel ratio, the steam to fuel ratio and the coal flowrate seem to be the main parameters


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affecting the reactor performance during the sensitivity analysis. The solid phase conversion depends on the air to fuel ratio while the gas

composition is mainly affected by the steam to fuel ratio. A double effect of steam injection is observed: at short values of the steam to fuel ratio the hydrogen content of the syngas increases while at high values the dilution effect prevails.

A high solid flowrate is found to cool off the reactor and bed extinction is observed for a solid flowrate equal to 0.006 Kg/s.

Differently from the biomass case the peak is less prominent and a wider high temperature region is present is the gasifier when the initial bed height is increased; a wider char heating section is formed in the gasifier in this case.

The higher is this parameter the longer is the heating section at steady state condition, pyrolysis occurs indeed at lower temperature when the initial bed is higher but devolatilization zone has almost the same length.

Further work is needed for this model. Model validation and calibration should be at this point carried out. A different calculation method should also be used in order to avoid stability problems and obtain a faster convergence since the time step in this case is ten times lower than in the biomass case and calculation time is in the order of 20 hours.

5.6 References





4. Qader S.A., Natural gas substitutes from coal and oil , 1985, Amsterdam, Elsevier, 1985


6. Yoon H., Wei J., Denn M. M., Transient behavior of moving-bed coal


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gasification reactors, AIChE Journal, 25, 1979, 429-439

Section II

Gas clean up: tar removal

Chapter

6

Iron material exploitation for biomass

tar removal: an experimental study

Within the general description of the gasification process is underlined the importance of a gas cleaning section. As previously discussed one of the most important problems in the case of biomass gasification is the tar content of the gas product because of its tendency to condense and react to give soot which can block filters and originate problems on downstream equipment. During a period of six months spent at KTH-Chemical technology division an experimental study on the effect of iron based catalyst for tar removal was conducted. The problem of tar removal is here discussed and results concerning tar removal and the effect of iron catalysts on syngas composition are presented in this section. Data is covered by secrecy agreements and just some results can be presented in this work.

CHAPTER 6 IRON MATERIAL EXPLOITATION FOR BIOMASS TAR REMOVAL

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6.1 Introduction

Addressing the topic of tar removal a unique and unequivocal definition of tar is needed. Many definitions are proposed in the literature. According to a dictionary available on the web [http://www.thefreedictionary.com] tar is “A dark, oily, viscous material, consisting mainly of hydrocarbons, produced by the destructive

distillation of organic substances such as wood, coal, or peat”. According to a more reliable source, a comprehensive work on tar nature

formation and conversion [15], tar is defined as “The organics, produced under thermal or partial-oxidation regimes of any organic material and are generally assumed to be largely aromatic.”

Tar is a mixture of different organic compounds connected to the thermal decomposition of organic matter. However, a detailed tar composition and its properties are still necessary. According to the EU/IEA/US-Doe meeting held in Brussel in 1998, dealing with the definition of a tar measurement protocol, tar is defined as “all organic contaminants with a molecular weight larger than benzene” [14]. This definition leads to a large variety of compounds so that many mixtures of organic molecules originated from thermo-chemical decomposition of organic matter are included.

It is nowadays well understood that tar compounds are a function of type of feed, reactor and operating conditions. The main parameters influencing the different processes are thermal environment and residence time. Figure 45 shows the relationship between the yield of tar and the reaction temperature as reported in [15] In Table 11 are shown classes of chemical components in pyrolysis products and gasifier tars from various processes based on GC/MS analysis [15]. Evans and Milne suggested that a systematic approach in the classification of the pyrolysis products can be used to compare products from the various reactors [8, 9, 15]. Four major product classes as a function of temperature of formation are indicated in their work:

� Primary products: characterized by cellulose-derived products such as levoglucosan, hydroxyacetaldehyde, and furfurals; analogous hemicellulose-derived products; lignin-derived methoxyphenols;

� Secondary products: characterized by phenolics and olefins; � Alkyl tertiary products: methyl derivatives of aromatics, such as

methyl acenaphthylene,methylnaphthalene, toluene, and indene;

SECTION II GAS CLEAN UP: TAR REMOVAL

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� Condensed tertiary products: the PAH series without substituent (benzene, naphthalene, acenaphthylene, anthracene/phenanthrene, pyrene).

Table 11: tar compounds as a function of process and temperature

Conventional Flash Pyrolysis (450°C–

500°C)

High-Temperature Flash Pyrolysis (600°C –650°C)

Conventional Steam Gasification (700°C –800°C)

High-Temperature Steam Gasification (900°C –1000°C)

Acids Aldehydes Ketones Furans

Alcohols Complex Oxygenates

Phenols Guaiacols Syringols

Complex Phenols

Benzenes Phenols

Catechols Naphthalenes

Biphenyls Phenanthrenes Benzofurans

Benzaldehydes

Naphthalenes Acenaphthylenes

Fluorenes Phenanthrenes Benzaldehydes

Phenols Naphthofurans Benzanthracen

es

Naphthalene* Acenaphthylene Phenanthrene Fluoranthene

Pyrene Acephenanthrylene Benzanthracenes

Benzopyrenes 226 MW PAHs 276 MW PAHs

* At the highest severity, naphthalenes such as methyl naphthalene are stripped to simple naphthalene

Figure 45: tar yield as a function of reaction temperature.


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As reported in [15] primary and tertiary products are mutually exclusive. Primary products are destroyed before that tertiary products appear: primary and tertiary compounds in the same tar sample would indicate non uniform conditions inside the reactor. It is usually believed that in a high temperature environment tar decomposes to give CO and H2, this is true with primary product cracking but this is not true for the condensed tertiary products. The tertiary compounds are usually more refractory and the subsequent cleaning section is affected by the decision to run gasification at more severe conditions. The removal of such refractory tar could be more difficult than that obtained from a mild gasification process even if a larger amount of tar is produced in the latter case.

Table 12: tar classification according to Devi et al.[5].

Class name Property Representative compounds

GC-undetactable Very heavy tars, cannot be detected by

GC None

Heterocyclic Tars containing hetero atoms; highly

water soluble compounds

Pyridine Phenol Cresols

Quinoline Isoquinoline

Dibenzophenol

Light aromatic Usually light hydrocarbons with single ring; do not pose a problem regarding

condensability and solubility

Toluene Ethylbenzene

XylSnes styrene

Light polyaromatic Two and three ring compounds;

condense at low temperature even at very low concentration

Indene Naphthalene

Methylnaphthalene Biphenyl

Acenaphthalene Fluorene

Phenanthrene Anthracene

Heavy polyaromatic Larger than three-rings, these components condense at high

temperatures at low concentrations

Fluoranthene Pyrene

Chrysene Perylene

Coronene


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Another classification is given in [4, 18], all the tar species are there grouped in six lumps, respectively called benzene, 1-ring compound, naphthalene, 2-ring compounds, 3- & 4- ring compounds and phenolic compounds. The same authors divided tar compounds in two groups: ‘easy to destroy’ and ‘hard to destroy’ depending on reactivity.

One important classification based on solubility and condensability of different tar compounds has been developed in cooperation with Energy research Center of The Netherlands (ECN), Toegepast Natuurwetenschappelijk Onderzoek (TNO) and University of Twente (UT) within the framework of the project ‘Primary measures for the inhibition /reduction of tars in biomass fuelled fluidised-bed gasifiers’, funded by the Dutch Agency for Research in Sustainable Energy (SDE) (see Table 12) [5].

Each type of gasifier, fed with the same feedstock, gives off tar with a characteristic chemical profile and this is true even within a specific family of gasifiers. These differences result from varying geometries, configurations, and temperature profiles, residence times, and bed materials.

6.2 Tar sampling and analysis

Tar analysis and characterization has received a wide interest among the gasification world and data on tar composition are widely available in the literature. Nevertheless, a problem comes from the heaviest tar compounds because they can condense on the sample filter or create soot particles in the sampling probe. As reported in [15] it is difficult to compare results from different works, moreover because some of the heaviest compounds are insoluble in certain solvents or seem to polymerize on the filter paper to form insoluble soot particles. The soot forming reactions are enhanced by the high temperature, sampling at low temperatures should be recommended, but this is in conflict with the need to sample at temperatures high enough (e.g., 400°C) to avoid tar condensation. A detailed description of the sampling method is therefore crucial while presenting data from tar analysis or during the comparison of many of them.

Two main alternatives for tar collection based on condensation or on solid phase adsorption exist. While the first route is based on condensation of tar in a solvent, solvent evaporation and gravimetric analysis of the sample, during the second the gas is passed through a solid phase that adsorbs tar compounds.


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Hydrocarbons are then extracted from the solid and analyzed in a GC or MS. During the meeting held in Brussels the need of a tar sampling protocol was

expressed. In 2002, as the result of a collaboration between many European countries a tar guideline was presented [14]. According to the guideline the tar and particle sampling system consists of a heated probe, a heated particle filter, a condenser and a series of impinger bottles or alternative equipment such as the Petterson column containing a solvent for tar absorption. The latter are placed in a temperature controlled bath. The gas is sampled for a well defined period through the sampling line and filter. Process pressure or a pump keeps the right flow rate for tar sample. In order to destroy the aerosols formed in the earlier bottles, a fit should be present in the impinger bottle.

Tar condensation should be avoided, sampling lines including the filter are therefore heated or a good insulation has to be provided. The temperature choice influences the thermal decomposition of organic compounds or its condensation, the selection of the right temperature is very important, some indications are given in the guideline. The tar collection occurs both by condensation and absorption in isopropanol which was found to be the most suitable solvent. The volume, temperature, pressure, and gas flow rate through the equipment are measured.

Immediately after sampling the solution has to be decanted into a dark storage bottle (in alternative the bottled sample must be stored in a dark place).

It is important to wash with the solvent all surfaces (including metal surfaces) contacting the gas at temperatures lower than the process temperature and to combine them with the actual sample.

The storage bottle has to be stored, tightly closed at a cool (< 5°C) temperature for analysis.

Sampling of tar and particles is usually performed simultaneously except for pressurized and/or large-scale gasifiers (>20 MWth) where a sampling strategy based on separate sampling of tar and particles is applied. Hence tar sampling is performed non-isokinetically for pressurized gases.

Non-isokinetic tar sampling is also practical in large-scale atmospheric gasifiers where the pipe diameter is large. Isokinetic sampling is also not required when only tar (and not particles) is sampled and the gas temperature under study at the sampling site exceeds 350°C. Such temperatures generally avoid tar condensation in the form of aerosols and/or droplets and also minimize adsorption of organic species onto particles.


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This system requires a large space, different tools and is time consuming. Some of these instructions are usually followed but many others methods such as cold trapping and solid phase adsorption (SPA) are commonly used in practice.

Cold trapping consists in tar condensation with the use of two condenser traps immersed in a cooling bath. According to the SPA method, developed at KTH [3], the sample is collected by adsorption and condensation at room temperature on an s.p.e. column containing 100 mg of amino phase through the use of a gastight syringe. Tar samples are taken by inserting the syringe needle into the process line via a rubber septum. By manually pulling product gases through the column, in 1 min a 100mL gas sample is collected. The column is capped in order to prevent contamination before tar desorption and analysis. An exhaustive description of the method is given in [3].

6.3 Tar removal

It is now clear that tar is an undesired component of the producer gas and that its removal is highly required.

Physical or chemical methods are used for tar removal. A brief summary of the different available or under development technologies is given in the following, in particular a description of the state of the art on catalytic cracking of tar will be presented.

6.3.1 Physical removal

Wet and dry systems are used to separate tar from the producer gas with physical methods, these employ well known technologies such as condensation, gas/liquid mixtures separations, droplet filtration and ceramic filters.

Cooling towers and Venturi scrubbers are the most important candidates for a wet tar removal section. The heaviest fractions of tar condense there while some droplets are still entrained by the gas flow. Moreover, tars require physical capture and agglomeration or coalescence more than simple cooling, unfortunately biomass-derived tars are very difficult to coalesce and a complex system is required to attain high removal efficiency. Tar separation efficiencies given in the literature range from 51% to 91% in a Venturi scrubber used to purify the producer gas from a countercurrent rice husk gasifier [15].


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Such efficiency is very low and the final tar content too much high for most of the end user applications of the syngas. Moreover, the amount of wastewater generated from a wet scrubbing system is very high and the contamination level of particulate and tar components is an important factor during the evaluation of these cleaning systems: the treatment cost is considerable and probably prohibitive in many cases.

Among the dry technologies, fabric, ceramic and metallic filters can remove near-dry condensing tar and particles from the gas product. Electrostatic precipitation is also an effective but costly way of removing tars but there is little experience on biomass-derived gasification products. Nevertheless, they can be only partially effective. If a liquid layer is formed on the surface of the filtering material it will cause mechanical problems and the efficiency of removal will decrease very quickly. However, a water cleanup system was tested in the Harboøre gasifier in Denmark [http://www.volund.dk] the system was found ineffective and abandoned and the scrubbing system was replaced by a gas cooler and a wet electrostatic precipitator.

A typical system for physical removal of tar should include a saturator (in order to cool and saturate the gas and allow the coalescence of particulates and tars in the next stage), a high-efficiency scrubber and a high-residence-time tower to allow the system to equilibrate. Tar levels down to 20-40 mg/m3 and particulate levels down to 20 mg/m3 can be achieved with such a system [15]. However, as previously discussed, these systems are expensive and create a waste disposal problem by generating large quantities of wastewater. Cooling of the product affects electrical efficiency (in case of the production of biomass to power production plant).

6.3.2 Chemical removal

Although physical methods are used to reduce the hydrocarbons level of the gas, they are very rarely effective so they can be used for a partial removal but further methods are required to achieve the low concentration admitted. Chemical methods are then used to break tar molecules. Once formed tar compounds are quite stable and high temperature or particular conditions are therefore needed in order to deplete their concentration. Cracking, steam reforming, dry reforming and carbon formation reactions are promoted during the syngas cleaning through the chemical route. This is achievable in two ways,


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exposing the gas to a high temperature environment or promoting these reactions in milder conditions with the use of a catalyst. Both the alternatives are discussed in the following.

6.3.3.1 Thermal reduction

Despite the possibility to reduce these mono-, di-, and polynuclear aromatic constituting tar to light gases under reducing conditions, a number of references about the requirement of high temperature (above 1000°C) or high residence times for thermal cracking are given in the literature. Thermal cracking is achievable in several ways: by increasing residence time after initial gasification, by direct contact with an independently heated hot surface or by partial oxidation by addition of air or oxygen. An overstatement of the freeboard section is for example required to achieve an efficient conversion of tar by thermal cracking at common operating conditions of a fluidized bed gasifier. In case of high temperature achieved by oxygen addiction or by contact with a hot surface the operative costs increase and the economical feasibility of the process is depleted. Nevertheless, the problem of soot formation and the refractory of heavier compounds formed at higher temperature should be avoided and the idea of a total removal of tar by thermal cracking does not seem suitable.

6.3.3.2 Steam Reforming

A work about gasification of biomass present in the literature indicates that tar produced in air gasification is more refractory than those produced in steam [18]. The idea of steam injection to the reactor in order to promote steam reforming reactions of tar should be proposed, nevertheless this could affect all the reactions involved in gasification and an adequate steam to fuel ratio comes out from the optimization of the overall gasification process. Because of this, the advantage of steam injection to the reactor on steam removal is limited and suitable just as a part of a global optimization of the reactor performance.

6.3.3.3 Catalytic cracking

Although the importance of physical processes and thermal cracking for tar removal is proven, their effectiveness is restricted and the promotion of cracking, steam reforming, dry reforming and carbon formation at lower


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temperatures appears more suitable. In order to achieve a significant conversion, the use of a catalyst is therefore needed in this case. A great variety of approaches have been tried in the catalytic conversion of tar.

The research in this field has involved primary methods and secondary methods. In the first case the catalyst is incorporated (for example with impregnation techniques) or mixed with the feed biomass (as a constituent of the bed material) to achieve the so-called catalytic gasification or pyrolysis while in the latter case raw gas is treated in a further catalytic reactor.

Two main classes of catalyst have been studied: non-metallic and metallic oxides. Many catalysts such as dolomite, olivine, limestone, Ni based catalysts, zeolites and char have been extensively examined and used in bed applications and in additional secondary reactors [1, 2, 5, 6, 10, 12, 13, 15-19, 21].

More recent work has included dual systems with catalysts such as dolomite serving as a guard bed for highly active catalysts like Ni-based reforming catalysts.

Results as presented in the papers are very often given briefly and incompletely, they are usually affected by the previously mentioned factor concerning the definition of tar, its composition and collection method.

Pilot-scale tests have shown that catalytic cracking of tars can be very effective. Tar conversion higher than 99% has been achieved using dolomite, nickel-based and other catalysts at elevated temperatures (typically 800-900°C) [15]. These tests have been performed using both fossil and renewable fuels. Most reported work used a second reactor. Some work has been carried out on incorporation of the catalyst in the primary reactor which has often been less successful than use of a second reactor [10]. Generally, gasification plants are equipped with a filtering system for particle removal and a separate tar conversion section.

Catalyst deactivation is generally not a problem with dolomite. An initial loss of activity is sometimes experienced as carbon compounds deposit on the catalyst, but these compounds gasify as the bed temperature rises and the catalyst is reactivated [15]. Metal catalysts tend to be more susceptible to contamination. Low hydrogen concentrations in the product gas reduce the catalytic activity of metal-based systems.

In previous works at KTH [16, 17] it has been confirmed that metallic iron has a pronounced ability to decompose tars produced during biomass gasification in a fluidized bed reactor increasing the gas yield. It has also be


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proven that unlike the use of other types of catalysts (e.g. dolomite) the employment of iron based catalyst increases the permanent gas yield but the methane content in the gasification gas is rather unaffected.

6.4 Experimental setup

This study was carried out at KTH - Chemical Technology Division during an internship period of six months within the gasification group. The main goal of the work was to test the capability to break down tar compounds of different iron based

catalysts. An atmospheric gasification unit available at KTH was used to carry out the experiments.

6.4.1 Material characterization

Catalyst properties are confidential and secrecy agreement exists between the research group and the catalyst providers, nevertheless, some basic information concerning the catalyst are presented in the following. Three types of iron material, supplied by Höganäs AB, were investigated (in the following they will be referred as catalyst A, B and C). The catalysts were characterized in terms of surface area (BET surface area was indicated by the producer), sulphur, nitrogen and oxygen content by Höganäs AB. In particular the oxygen content of the catalyst was monitored before and after running the gasification experiments, in order to see if the oxygen potential of the system was sufficiently low to keep the iron material at metal state. Some oxidation reactions were observed when a mixture of nitrogen and oxygen equal to the gas feed was fed to the reactor in absence of solid feed for a long time (see Figure 46). However, those oxygen levels are quite far from the real content of a regular run and analysis showed that the oxygen content of the fuel did not change before and after the catalyst use during conventional runs. A list of catalyst properties with reference to the catalyst type is shown in Table 14.


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Metallic iron was tested with real syngas coming from a gasification unit (see Figure 47, Figure 48 and paragraph 6.4.2 for a detailed description of the unit) fed with biomass. Feeding properties of two kinds of wood were initially tested, a woody mixture mainly constituted by red fir, and Swedish birch wood. The first was obtained from milling and sieving of wood chips, particles shape and size distribution was unsuitable in relation to the feeding system in this case especially because of the high ratio between length and width of the particles. Their use was therefore abandoned. Two types of Swedish birch wood were tested instead, with and without bark. Because of better feeding properties the latter was chosen between the two. The first step on feeding preparation consisted on wood sieving, a particle size of 1-2.5 mm was chosen. The moisture content of the fuel was analyzed. The analysis is based on the weight loss of biomass due to the exposure at 105° C for 24 hours.

An average moisture content of 5.9 wt% was found between the different experimental runs.

The bed material consists of Al2O3, with a particle size fraction ranging between 63µm and 125µm. Nitrogen is used as the main fluidizing agent and oxygen is mixed with nitrogen according to the actual experimental parameters.

Figure 46 catalyst oxidation due to high oxygen concentrations in the catalytic bed


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Table 13: ultimate analysis of birch wood used as the feedstock for the experiments.

Element Weight %

C 49.0

H 6.2

N 0.3

O 44.5

Table 14: properties of the three iron materials tested.

BET

m2/kg

Concentration %

N O S

A 285 0.000 0.801 0.001 B 281 0.013 0.762 0.005

C 181 0.000 0.367 0.002

6.4.2. Apparatus

Experiments were performed in an atmospheric, fluidized bed gasification system developed by KTH in cooperation with Termiska Processor i Studsvik AB (TPS) [10]. The system and its schematization are respectively given in Figure 47 and Figure 48.

It is provided by three electrical heaters plus an additional heater for fluidization gas preheating.

A very crucial point of the reactor is the distribution disk delimiting the reactor at the bottom, due to the ash content of the fuel and high temperature of the bed sintering problems and crack formations with a consequent loss of fluidization was many times observed. The possibility to run the gasifier until a temperature equal to 950°C exists.

A ceramic filter for soot and entrained bed material removal is present between the gasifier and the catalytic reactor. The filter is provided of an external heater in order to prevent condensation of tars, the same temperature as the fluidized bed is kept in this section. Temperature of the catalytic reactor is guaranteed by the use of an additional electrical heater, 900°C is the maximum operating temperature for this section. The filter and the catalytic fixed bed reactor have a length of 0.70 m and an inner diameter of 0.05 m.

The possibility to feed the fuel directly into the fluidized bed near the gas distributor plate or to the top of the gasifier exists. The first option was chosen


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during the experiments. The feeding section is constituted by a storage vessel and a screw feeder. In order ensure the equilibrium of pressure between the feeder and the reactor and thus allow a correct performance of the feeder a nitrogen flow (secondary nitrogen) to the storage vessel is provided. A more detailed experimental setup has been described by Vriesman [20].

6.4.3 Experimental performance

Each experiment consisted in four different steps: materials preparation, heating up of the gasifier, biomass feeding and steady state achievement and finally tar sampling and gas analysis. The bed material consisted of 350 g of Al2O3, biomass was sieved to the desired size and the vessel was filled up. Biomass consumption for every run was calculated as the difference between biomass put in the vessel and residual biomass after gasification. In order to keep the pressure equilibrium discussed in the previous paragraph the whole biomass needed for the experiment is introduced in the storage vessel before the experiment start and secondary nitrogen was fed to the reactor from the

Figure 47: schematization of the gasification system


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beginning. The catalyst container was filled with 150g of iron material. Electrical heaters were used to heat up the whole system. In previous experiments at KTH a start up procedure was developed, according to it slow heating is preferred and a constant nitrogen flow is fed to the reactor in order to avoid sintering problems at the distribution disk and in order to increase the heat transfer inside the bed ensuring a homogeneous temperature in the reactor. The gasifier was slowly heated up to the set up temperature. This operation lasted about three hours. Once that a uniform constant temperature was achieved inside the reactor, a defined engine frequency of the electrical motor was set and the screw was turned on. Oxygen was fed at the same time. About one hour of unsteady state was observed and just then the heart of the experiment could start, syngas was then analysed and tar collected. In order to ensure a constant temperature inside the reactor heaters were left on during the whole run. Gasification was allothermal.

As previously discussed the performance of the catalyst bed reactor was the aim of the study, the syngas composition and the tar content had to be kept as constant as possible during the experiment and between different runs, the possibility of every noisy caused by the gasifier performance avoided. After the experiment the char and bed material were collected, sieved and weighed and a material balance was calculated. The biomass feed was observed to be in the

Figure 48:experimental apparatus


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range of 3.95-4.25 g/min at the actual frequency. The bed temperature was set to 850°C for all the experiments. The flow rate of nitrogen (the fluidisation media) was 9.50 l/min and the oxygen flow rate was 0.85 l/min resulting in a lambda value of approximately 0.2. Temperature inside the catalytic bed was set to 700, 750, 800 and 850°C.

6.4.4 Gas and tar analysis

The gas sampling system is situated after the catalytic reactor. The composition of the gas was measured with a Shimadzu gas chromatograph equipped with a flame ionization detector (FID) and a thermal conductivity detector (TCD). Steam present in the raw gas is condensed before that the gas is fed to the gas chromatograph, and tar is removed. A system of impinger bottle and a washing column are used to reach this goal.

The SPA solid phase adsorption method as described in [3] was used for tar sampling and composition detection.

In order to examine the reliability of the test data each run was performed twice and three or four samples were collected for every temperature. The results from the different tests are in fairly good agreement. The data reported are average values of the two runs. A typical syngas composition before the catalyst is given in Table 15 while the tar content of the raw gas after the filter is presented in Figure 50.

Figure 49:gas cleanup system before the GC.


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6.5 Results and discussion

The tar content in the gas and the gas composition were measured at the inlet and the outlet of the catalytic bed. An unsteady state behaviour of the reactor was observed during the experiment. As shown in Figure 51 the methane and tar content of the gas before the catalyst bed decreases with time. The reason can be found on the char and ash accumulation inside the fluidized bed. It is indeed well known that they affect tar conversion during the gasification process. Moreover, this behaviour was observed for long lasting time experiments up to 3-4 hours while just small fluctuations were observed

Figure 50: tar composition before the catalytic bed.


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for short runs. In order to avoid the influence of this phenomenon on the results, long

lasting time experiments should be avoided. Data collected was analyzed and the tar conversion at every temperature for

every catalyst type were calculated as the percentage of converted tar (inlet – outlet) with respect to the tar content of the raw syngas. Results concerning tar removal and effect on syngas composition are presented in this section. Further data and results are covered by secrecy agreements and cannot be presented in this section.

6.5.1 Catalyst performance and tar conversion

In Figure 52 is presented the tar conversion as a function of bed temperature for the different catalysts. In accordance with theory the higher is temperature, the higher conversion is obtained in the catalytic bed. Moreover, at 700°C catalyst influence seems to be almost absent and tar leaves the reactor almost unconverted.

The highest conversion, equal to 80%, was achieved with catalyst A when the temperature of 850 °C was set in the bed. A conversion equal to 65% and 58% was respectively observed with catalyst B and C.

Figure 51: time variation of the tar content before the catalytic bed.

0

2

4

6

8

10

0

40

80

120

160

200

0 50 100 150 200 250

Met

hane

[%vo

l]

Tar

[µg/

100m

l]

Time [min]

Tar Methane


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The reason of this difference has to be found in the surface are of the three materials, catalyst A has indeed the higher surface area while catalyst C has the lower.

6.5.2. Influence on gas composition

One of the main objectives of the project was to find a catalyst able to break tar increasing the methane content of the gas or at least leaving it almost unchanged. The gas composition was monitored while the tar sample was taken and results show that it did not change substantially during the run. Small fluctuations in the gas composition at different temperatures were observed for all the iron materials. The variation of syngas composition with temperature and type of catalyst is shown in Table 15.

6.5.3 Catalyst Life

A real test on the catalyst life has not been performed yet, but many experiments were conducted with the same catalyst and no deactivation was observed in the catalyst for 20 h-on-stream tests on the operating conditions

0

20

40

60

80

100

650 700 750 800 850 900

Con

vers

ion

[%]

Temperature [°C]

A B C Without

Figure 52: catalyst performance


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previously presented.

Table 15: gas composition before and after the catalytic bed for the best performing catalyst

CH4 CO2 CO H2

Catalytic Bed Inlet [vol %]

8.15 19.94 44.79 26.20

Temperature °C Catalytic Bed Outlet [vol %]

700 8.35 21.20 42.11 26.81 750 8.24 21.79 42.83 26.73

800 8.39 20.86 43.58 26.60

850 7.86 19.26 46.05 26.68

6.6 Conclusions

Part of a project aimed at demonstrate the possibility to break tar produced during gasification of biomass but, at the same time, keep the methane content of the gas is here presented.

Three types of iron based materials provided by Höganäs AB were tested with real syngas. A fluidized bed gasification system equipped with a ceramic filter and a further fixed bed reactor for tar conversion was used to achieve this goal. The composition of the gas was measured with a Shimadzu gas chromatograph equipped with a FID and a TCD. Tar was sampled and analyzed with a solid phase adsorption method.

Samples were tested at three different temperatures, respectively 700° C, 750°C, 800°C and 850 °C. The three iron based materials showed an increasing activity with temperature, all of them seem to be suitable for tar removal at a bed temperature equal to 850 °C. Type A was the more effective since the higher conversion, equal to 80%, was achieved in this case. This result is also in accordance with the difference in surface area between the catalysts.

The gas composition was also monitored and just small fluctuations were observed for the three materials at different operating conditions.

The same sample was tested many times, no deactivation study was performed, but for operating conditions previously described any deactivation was observed during the 20 h on stream tests.


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6.7 References

1. Binlin Dou, Jinsheng Gao, Xingzhong Sha, Seung Wook Baek, Catalytic cracking of tar component from high-temperature fuel gas, Applied Thermal Engineering, 23, 2003, 2229–2239.

2. Binlin Dou, Weiguo Pan, Jianxing Ren, Bingbing Chen a, Jungho Hwang, Tae-U Yu, Removal of tar component over cracking catalysts from high temperature fuel gas, Energy Conversion and Management, 49, 2008, 2247–2253.

3. Brage C., Yu Q., Chen G., Sjöström K., Use of amino phase adsorbent for biomass tar sampling and separation, Fuel, 76, 1997, 137-142.

4. Corella J, Caballero MA, Aznar MP. A 6-lump model for the kinetics of the catalytic tar removal in biomass gasification, Proceedings of the first world conference on biomass for energy and industry, 2000.

5. Devi, L., Ptasinski, K.J., Janssen, J.J.G., van Paasen, S.V.B., Bergman, P.C.A., Kiel, J.H.A., Catalytic decomposition of biomass tars: use of dolomite and untreated olivine, Renewable Energy, 30, 2005, 565-587.

6. Ekström C., Lindman N., Pettersson R., Catalytic conversion of tars, carbon black and methane from pyrolysis/gasification of biomass, Fundamentals of Thermochemical Biomass Conversion, Colorado USA 1982

7. El-Rub Z. A., Bramer E.A., Brem G., Experimental comparison of biomass chars with other catalysts for tar reduction, Fuel, 87, 2008, 2243–2252.

8. Evans, R.J., and Milne, T.A., Chemistry of tar formation and maturation in the thermochemical conversion of biomass. In: Developments in thermochemical biomass conversion 2 (1987), Bridgwater, A.V., and D.G.B. Boocock, D.G.B. (Eds.), Blackie Academic & Professional, London, pp. 803- 816.

9. Evans, R.J., and Milne, T.A., Molecular characterization of the pyrolysis of

biomass. 2. Applications. Energy and Fuels, vol. 1, 1987. 10. F. Miccio, et al., Biomass gasification in a catalytic fluidized reactor with beds

of different materials, Chem. Eng. J. (2009), doi:10.1016/j.cej.2009.04.002

11. Heginuz E., Gregertsen B., Sorvari, Vriesman, Sjöström K., Studies on solid fuel pyrolysis and fluidisation in a laboratory-scale, atmospheric, fluidised

bed reactor, Finnish-Swedish Flame Days, Proceedings, Naantali, Finland, 1996.


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12. L. Ma, H. Verelst, G.V. Baron, Integrated high temperature gas cleaning: Tar removal in biomass gasification with a catalytic filter, Catalysis Today, 105, 2005, 729–734.

13. Li C., Suzuki K., Tar property, analysis, reforming mechanism and model for biomass gasification – An overview, Renewable & Sustainable Energy Reviews, 13, 2009, 594-604.

14. Maniatis K., Beenackers A. A. C. M., Tar protocols. IEA Bioenergy gasification task, Biomass and Bioenergy, 18, 2004, 1-4.

15. Milne, T.A., Abatzoglou, N.,Evans, R.J., Biomass Gasifier “Tars”: their nature, formation and conversion. Report NREL/TP-570-25357, NREL, Golden, Colorado, USA, 1998

16. Nordgreen T., Liliedahl T., Sjöström K., Elemental iron as a tar breakdown catalyst in conjunction with atmospheric fluidised bed gasification of

biomass: a thermodynamic study, Energy and Fuels 20, 2006, 890-895. 17. Nordgreen T., Liliedahl T., Sjöström K., Metallic iron as a tar breakdown

catalyst related to atmospheric, fluidised bed gasification of biomass, Fuel, 85, 2006, 689-694.

18. Perez P, Aznar PM, Caballero MA, Gil J, Martin JA, Corella J. Hot gas cleaning and upgrading with a calcined dolomite located downstream a biomass

fluidised bed gasifier operating with steam–oxygen mixtures, Energy Fuels, 11, 1997, 1194–1203.

19. Sjöström K., Taralas G., Liinanki L., Sala Dolomite-Catalysed conversion of tar from biomass pyrolysis, Research in Thermochemical Biomass Conversion, Phoenix, Arizona, USA, 1988.

20. Vriesman P., Nitrogen conversion in a biomass fuelled atmospheric fluidized bed gasifier, Licentiate of engineering thesis, KTH, 2001.

21. Zhang R., Robert Brown C., Suby A., Cummer K., Catalytic destruction of tar in biomass derived producer gas, Energy Conversion and Management, 45, 2004, 995–1014.

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Conclusions

This work deals with gasification of coal and biomass. Two aspects of the process are investigated: Section I deals with a mathematical model of an updraft reactor; Section II deals with the problem of tar removal through the exploitation of iron material as a catalyst for tar cracking.

The model presented in Section I has been used to study the behavior of an updraft fixed bed gasifier. Two case studies are analyzed: a birch wood feedstock and a Pittsburg seam n°8 coal seam feedstock. The dynamics of the system, the steady state behavior and a sensitivity analysis are reported in this work. In particular the dynamics is analyzed with reference to the bed behavior in a long time investigation. Scarce information about this topic was before present in the literature.

In every case study analyzed it is observed that, since a difference in the spatial velocity of the two phases along the gasifier exists, every change in the solid is gathered from the gas phase which quickly adapts to the new conditions. In order to observe the whole dynamics of the system, the behavior of the bed should be monitored.

The air to fuel ratio, the steam to fuel ratio and thermal conditions of the reactor are detected as the key parameters for the process.

A bed extinction is observed when a high amount of steam is injected to the reactor. The dilution effect of steam seems to be the reason of bed extinction.

Some temperature oscillations, due to the numerical method chosen for the solution of the equations, are observed in the combustion region in the case of a biomass feedstock.

At steady state a temperature peak is present in the combustion region while temperature is very low in the rest of the reactor. For this reason the fuel heating section and the drying and the pyrolysis section are longer than expected. This is a demonstration that temperature drives the system and a proper description of the reactor cannot abstract from a good description of the thermal conditions within the bed. Every experimental investigation of the system should be besides carried out with a good temperature monitoring.

The air to fuel ratio is strictly related to the solid phase conversion while the steam to fuel ratio mainly affects the gas composition. A double effect of steam injection is observed: at short values of the steam to fuel ratio the hydrogen

144

content of the syngas increases; at high values the dilution effect prevails. Further work is needed for this model. Model validation and calibration are

needed. The set of reactions can be improved, especially the improvement of the kinetic description is an interesting issue. Besides, the fate of Sulphur and nitrogen inside the reactor should be investigated. A different calculation method should also be used in order to avoid every stability problem.

In section II is presented an experimental investigation carried out during a collaboration with KTH (Sweden).

Three types of iron based materials provided by Höganäs AB were tested in a laboratory atmospheric gasification system.

Samples were respectively tested at 700° C, 750°C, 800°C and 850 °C. Their effectiveness on tar removal was confirmed. Moreover, they showed an increasing activity with temperature. One of them was the more effective since the higher conversion, equal to 80% was achieved in this case at 850°C. This result is also in accordance with the difference in surface area between the different catalysts.

Iron material was also found to keep almost a constant gas composition exactly as it was asked.

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Acknowledgements

I would like to thank Prof. Giampaolo Mura for his support during the last

years but most of all for teaching me the relevant importance of collaboration and knowledge sharing within the research field. Thanks to him I could also understand many aspects of life that I had ignored before.

Thanks to my colleagues at the Chemical Engineering and Material Science Department for the nice moments we spent together in our room. Special thanks to Andrea and his patience: he supported me in the last year and he listened to all my complaints.

The Chemical Technology division of KTH is kindly acknowledged, especially the gasification group and his leader Prof. Krister Sjöström for his support during my stay in Sweden. Special thanks to Vera Nemanova for her help and availability at my arrival to Stockholm, the wonderful time spent together at the lab and for being a good friend.

My Swedish, Italians, German and Bolivian friends are also acknowledged for the wonderful time I had in Sweden.

All the people who collaborated to this work are also kindly acknowledged. Because of many circumstances I was very moody, I thought life out, I

messed up and smarten up my mind during the three years. I should thank Cristiana, Ivana, Walter, Gianluca, Stefania, Laura, Alessandro

and Giovanni for supporting me in this period, thanks for staying with me even when I was not that kind, thanks for making me smile, thanks for believing in me, thanks for understanding me because I know that It was not always that simple to be there with me.

Above all I am grateful to my parents. They remind me who I am and where I belong to. Thanks for giving me the possibility to be what I am and for being there even when they do not agree with me or when they do not understand my purpose.

Everything that I am descends from their confidence, the freedom they gave me and the open mind they built while bringing me up.

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Communication to international and

national conferences related to this work

1. M. Brundu, A. Lallai and G. Mura – A mathematical model for coal and

biomass gasification in a fixed bed reactor – CCT 2007, Third Int. Conf. on Clean Coal Techn., Cagliari, May 15 – 17, 2007

2. Mura G., Lallai A., Brundu M. – Mathematical modelling of an updraft coal

and biomass gasifier – Proceedings PRES 2008, Praha (CZ), 24-28 Aug. 2008, 11 pages

3. M. Brundu, A. Lallai, G. Mura – Gassificazione di carbone e biomassa –

Convegno GRICU 2008, Le Castella (KR), 14-17 settembre, Atti pag. 589 – 594

4. Brundu M., Mura G., – Transient behaviour of a fixed bed counter current

gasifier: one dimensional modelling, – Proceedings of the 4th Int. Conf. on Clean Coal Techn., Dresden, May 2009

5. Brundu M., Lallai A., Mura G. – A simplified model of a fixed bed counter

current gasifier – Proceedings of the 4th Int. Conf. on Clean Coal Techn., Dresden, May 2009

6. Brundu M., Mura G. – A mathematical model for an updraft biomass gasifier –

Proceedings of the 17th European biomass conference and exhibition, Hamburg, June 2009

7. V. Nemanova, M. Brundu, T. Nordgreen, K. Sjöström, – Biomass

gasification in an atmospheric fluidised bed: probability to employ metallic iron as a tar reduction catalyst – Proceedings of the 17th European biomass conference and exhibition, Hamburg, June 2009.

gasification of solid fuels: countercurrent reactor...

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