Design of a Fluidized Bed Reactor for Biomass Pyrolysis

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Design of A Fluidized Bed Reactor For Biomass Pyrolysis

A thesis submitted to the

Graduate School

Of the University of Cincinnati

In partial fulfillment of the

Requirements for the degree of

Master of Science

In the department of Mechanical Engineering of the

College of Engineering and Applied Science

Alaba Olumide Bamido

B.Sc. University of Lagos

Committee Chair: Jude Iroh

June 2018 Abstract

Global warming is caused by the buildup of heat trapping gases such as carbon dioxide in the atmosphere. These gases retain heat from the sun, which makes the environment warmer. The use of biomass as a fuel fossil can help reduce the amount of carbon dioxide in the atmosphere. When fossil fuels such as coal, gas and crude oil, which are locked underground, are burnt in oxygen, they release an enormous amount of carbon dioxide to the atmosphere. Aside burning of fossil fuel, decaying plant residues also contribute to the increase in atmospheric carbon dioxide. Biomass derived fuels minimize dependence on underground fossil fuel thereby preventing excessive amount of carbon dioxide to be released to the atmosphere.

Biomass is regarded as a renewable fuel because it can be replenished over a short period of time and is an active component of the carbon cycle. During plants (biomass) life, significant amount of CO2 is absorbed through photosynthesis. At death, these CO2 is released back to the atmosphere therefore ensuring no new carbon dioxide is produced. Pyrolysis of biomass offers many environmental benefits, most importantly, the reduction of CO2 in the atmosphere. Aside the reduction of carbon dioxide in the atmosphere, the biochar obtained from the pyrolysis of biomass can be used for soil amendment. Biochar in the soil improves soil fertility, reduces erosion and leaching of soil nutrients.

Conventional ways of processing biomass involves a centralized system where all the biomass in a region are gathered at a specific location and transported to a commercial pyrolyzing plant where they are processed. A major disadvantage of this method is the use of a large area of land for construction of the biomass pyrolysis plant. It is also strenuous, bulky, immobile, and requires adequate planning and logistics for successful operation.

This research project solves many of the above-mentioned issues. A small size and mobile fluidized bed reactor is designed and validated in this work. The design outlined only shows a first iteration of the overall design process. Parameters such as design pressure and temperature were chosen from arbitrary but reasonable estimates. The advantages of a fluidized bed pyrolysis reactor over other reactor types are: simple design, easy to repair when faulty, uniform temperature of particles, efficient mixing of particles and relatively cheap manufacturing cost. Solidworks was used to design the reactor geometry and ANSYS FLUENT was used for the thermochemical conversion analysis. ANSYS STATIC STRUCTURAL was also used for the Thermal and structural analysis. In addition, a cost estimate for the design is also presented.

Acknowledgement

I want to thank God for giving me the strength and grace to start and complete this work.

My appreciation goes to the faculties and staffs of University of Cincinnati. The faculties in Mechanical & Materials Engineering, and Aerospace Engineering provided me with a solid understanding of engineering principles which was very valuable to enable me complete this research project.

Special thanks to my advisors, Dr. Ronald Huston and Dr. Jude Iroh for their mentorship and advice during the course of this research project. Big thanks to Dr. Ronald Huston for making it possible for me to pursue a Master’s of Science Degree at the University of Cincinnati.

My appreciation also goes to faculties and staffs of Chemical and Environmental Engineering (ChEE) for their care during my studies. They were like a family to me even when I was very far away from home.

Finally, I would like to thank my family members, my Dad, Mum, my five siblings and relatives for their love and support during my studies. And also thanks to a whole lot of friends who made it possible (directly or indirectly) for me to complete this work.

A big thanks to you all.

Table of Content

Abstract……………………………………………………………………………………….……..1

Acknowledgement…………………………………………………………………………………..2

List of Figures…………………………………………………………………………………….....5

List of Tables...…………………………………………………………………………………...…7

Nomenclature...…………………………………………………………………………………...…8

Chapter 1: Introduction……………………………..……………………………………………...10 1.1: Problem Statement………………………………………………………….…………10 1.2: Research Objective……………………………………………………………...…….10

Chapter 2:Literature Survey…………………………………………………………………...…...12 2.1: Constituents of Biomass...…………………………………………...…………...…..12 2.2: Pyrolysis……….….……………………………………………………………...…..14

Chapter 3: Types of Pyrolyzers for Biomass Pyrolysis…………………………………………....18 3.1: Fast Pyrolysis Reactors…...………………………………………………………...…18 3.2: Slow Pyrolysis Reactors…………………………………………………………...….23

Chapter 4: Fluidization and Design of a Fluidized Bed Reactor………………………………..…27 4.1: Fluidization theory…………………………………………………………………….27 4.2: Geldart Classification of Particles…………………………………………………….29 4.3: Design of a Fluidized Bed Reactor……………………………………………………30

Chapter 5: FEA Analysis…………………………..…………………………………………....…50 5.1: Final Reactor Geometry……………………………………………………………….50 5.2: Meshing……………………………………………..…………...……………………51 5.3: Boundary Conditions……………….………………………………...……………….51 5.4: Results………………………….………………………………………...……………51

5.5: Results Verification/Validation……………………………………………...………..63

Chapter 6: Reaction Kinetics of Biomass Pyrolysis, CFD Modelling and Product Prediction…....64 6.1: Reaction Kinetics of Biomass Pyrolysis…………………………………………...….64 6.2: CFD Modelling and Product Prediction of Biomass Pyrolysis…………………….....65 6.3: Results………………………………………………………………………………....72 6.4: Verification and Validation of Results……………………………………………..…77

Chapter 7: Cost Analysis……………………………………...………………………………...…79

Chapter 8: Conclusion………………………….………..………………………………………....80

Chapter 9: Future Directions……………………………...…………………………………...…...80

References……………………………..………………………………………………………...…81

Appendix...……………………………………………………………...……………………...…..85

List of Figures

Fig 2.1: Cellulose structure………………………………………………………………………13

Fig 2.2: Structure of Hemicelluloses………………………………...…………………………..13

Fig 2.3: Some structural units of lignin………………………………………………………….14

Fig 3.1: Bubbling fluidized bed pyrolyzer……………………………………………………….19

Fig 3.2: Circulating Fluidized Bed Reactor………………………………………………...……20

Fig 3.3: Rotating cone pyrolyzer reactor………………………………………….……………..21

Fig 3.4: Ablative pyrolyzer………………………………………………………………………21

Fig 3.5: Twin screw pyrolysis process…………………………………………...………………22

Fig 3.6: Ultra-rapid pyrolyzer……………………………………………………………………23

Fig 3.7: Beehive oven as a fixed pyrolyzer for charcoal production…………………………….24

Fig 3.8: Heated Kiln pyrolysis process…………….…………………………………………….25

Fig 3.9: Screw/Auger pyrolysis process…………………………………………………………26

Fig 4.1: Schematic representation of the different regimes of fluidized bed by Kunni and Levenspiel……...………………………………...………………………………………………28

Fig. 4.2: Diagram of Geldart particle classification………………..…………………………….30

Fig 4.3: Different types of distributors…………………………………………………………..34

Fig 4.4: Nozzle Jet Configurations………………………………………………………………39

Fig 4.5: Fig 4.5: Distributor plate and nozzles…………………………………………………...46

Fig 5.1: Fluidized Bed Reactor…………………………………………………………………..50

Fig 5.2: Geometry of reactor……………………………………………………………………..52

Fig 5.3: Mesh of Reactor Geometry…………………………………………………………..…53

Fig 5.4: Reactor Geometry showing boundary condition (Pressure, Force and Fixed support)...55

Fig 5.5: Section of Reactor geometry showing boundary condition (Pressure, Force and Fixed Support)…………………………………………………………………………………………..56

Fig 5.6: Reactor Geometry showing thermal boundary condition……………………………….57

Fig 5.7: Total Deformation of the Reactor under operation…………………………………….58

Fig 5.8: Section through the reactor showing Total Deformation………………………………59

Fig 5.9: Equivalent (von-Mises) stress simulation result……………………………………….60

Fig 5.10: Section through the reactor showing Equivalent Stress Simulation result……………61

Fig 5.11: Section through the reactor showing Equivalent Elastic Strain Simulation result……62

Fig 6.1: One stage multiple reaction……………………………………………………………..64

Fig 6.2: Two stage semi-global reaction…………………………………………………………65

Fig 6.3: Geometry of a 2D section of the reactor shell designed with ANSYS R19.0 DesignModeler………………………………………………………...…………………………63

Fig 6.4: Inlet Boundary Condition (on the sides of nozzles)…………………………………….67

Fig 6.5: Wall boundary condition………………………………………………………………..68

Fig 6.6: Outlet boundary condition………………………………………………………………69

Fig 6.7: Geometry with meshing applied………………………………………………………..70

Fig 6.8a: Contour of Total Pressure acting in the reactor due to fluidization……………………72

Fig 6.8b: Magnified view of Total Pressure contour…………………………………………….73

Fig 6:9: Contour of velocity magnitude of particles in the reactor………………………………74

Fig 6.10: Contour of y-velocity of the particles in the reactor…………………………………...75

Fig 6.11: Contour of Static Temperature in the reactor………………………………………….76

Fig 6.12: Plot showing the pyrolysis product formation as a function of time…………………..77

List of Tables

Table 4.1: Material properties and justification………………………………………………….30

Table 4.2: Reactor parameters…………...………………………………………………………31

Table 4.3: Reactor Design Analysis Parameters…………………………………………………40

Table 4.4: Plenum Design Criteria and Formula as proposed by Litz…………………………...49

Table 5.1: Reactor Geometry Meshing Details…………………………………………………..54

Table 6.1: Biomass pyrolysis kinetic properties by Thurner et al……………………………….71

Table 6.2: Biomass pyrolysis result……………………………………………………………...77

Table 6.3: Comparism of research result with known experimental results……………………..78

Table 7.1: Engineering Bill of Materials………………………………………………………...79

Nomenclature

Ar – Archimedes number

퐶퐷 – Orifice coefficient

D – Diameter of reactor

푑푒푛푡푟푦 – diameter of inlet pipe to the plenum

푑푛 – diameter of nozzle

푑표 – diameter of nozzle orifice

푑푝 – diameter of particle

E – Joint Efficiency

휌푔 – density of fluidizing gas

휌푝 – density of particles in the reactor

휌푏 – density of biomass in the reactor

푓푠 – Factor of safety g or 푔 – acceleration due to gravity

퐻푚푓 – Bed height at minimum fluidization

ℎ푠 – height of sand

ℎ푏 – height of biomass

ℎ푝 – height of particles

ℎ푝푙푒푛푢푚 – height (vertical) of plenum

퐿푑표푤푛 – length of downward directed jet

퐿ℎ표푟 – horizontal length of horizontally directed jet

푙푗 – jet height/length where j may be upward, horizontal or downward

푙푚 – minimum nozzle length

퐿푢푝 – vertical length of upward directed gas jet

푚푠 – mass of sand in reactor

푁 – number of orifices per unit area

푛표 – number of orifices per nozzle

푛 – number of nozzles

P – Design Pressure

∆푃퐵 – Pressure drop across the bed

∆푃퐷 – Distributor plate pressure drop

푅푒푝 – Reynolds number of particles

푅푒푝,푚푓 – Reynolds number of particles at point of minimum fluidization

푟 – radius of distributor plate

S – Allowable stress

T – Design Temperature

푡푤 – Wall thickness t or 푡푝 – thickness of distributor plate

푢 – superficial velocity

푢푚푠 – minimum slugging velocity

푢푚푓 – minimum fluidizing velocity

푢표 – velocity of the gas at the orifice

푢푡 – terminal velocity

푣푏 – volume of biomass in reactor 푣 – poisson ratio

Chapter 1: Introduction

1.1 Problem Statement

The increase in atmospheric CO2 is a major environmental problem affecting our planet today. Its effect can be felt in several aspects of our daily lives. The most prominent effect of this adverse build up of CO2 gas is global warming. CO2 being a heat trapping gas absorbs heat from the sun which makes our planet warmer. This in turn causes the melting of ice caps leading to increase in ocean/sea levels thereby flooding coastal cities and communities.

There are several causes of CO2 emissions in our atmosphere. These include burning of fossil fuels, automotive combustion, forest fire, animal respiration, decaying of plants and animal remains etc. Of these sources, combustion of fossil fuel accounts for the greatest supply of atmospheric carbon dioxide. The fossil fuel underneath the earth does not pose any threat to our planet. However, when these fuels are drilled and burnt, a great amount of carbon dioxide is released to our atmosphere. The process of fossil fuel formation from plant and animal remains usually occur over millions of years and is not easily replenished which implies that more CO2 is released to the atmosphere than fossil fuel is formed. Therefore, a carbon negative process is required to eliminate some of the carbon dioxide in the atmosphere.

During the lifetime of a plant, it absorbs carbon dioxide for manufacturing its food through the process of photosynthesis, in the presence of sunlight. At death or during decay of plant residue, a large amount of the carbon dioxide absorbed by the plant is released back to the atmosphere. Several factors are responsible for plant death or decay of plant residue. These include deforestation, seasonal shedding of plant leaves, drought etc. In this research work, a process is developed to minimize the release of carbon dioxide from decaying plant residues.

1.2 Research Objective

The objective of this research work is to design a pyrolysis reactor to convert the decaying plant residues into other useful energy forms that pose less threat to the environment. Pyrolysis of biomass is the thermal decomposition of biomass in the absence of oxygen or very little oxygen. The pyrolytic conversion of biomass gives three major products (biochar, biooil and biogas). The biochar can be released to the soil to improve soil amendment, prevent erosion and leaching of plant nutrients while the biooil and biochar can serve as alternative energy sources thereby reducing the dependence on fossil fuel. A fluidized bed reactor was designed because of the following advantages it has over other reactors:

1. Relatively simple and less expensive design 2. Ease of repair 3. Efficient mixing of particles in the reactor

4. Uniform temperature distribution in the reactor

The parameters used in the design were arbitrarily but carefully chosen within a reasonable limit. The design carried out in this research is the first design iteration in the overall design process. Several commercial fluidized bed pyrolysis reactor exist today. These design are fixed to a particular location and constructed over a large area of land. The design in this research differs from existing design because of its compact and smaller size and the ability to be moved from one location to another. The design proposed is validated and the result yield of the pyrolysis is compared to known experimental values.

Chapter 2: Literature Review

Biomass generally refers to renewable plant organic matter developed by the process of photosynthesis [1]. It also includes liquids and gases recovered from the decomposition of biodegradable and non-fossilized organic materials. It is a renewable energy source because it replenishes within a very short time. As a renewable energy, biomass is formed by the interaction of carbon dioxide, air, soil, water and sunlight with plants and animals. When an organism dies, microorganisms break down the remnants which eventually lead to the release of carbon dioxide stored during the lifetime of the organism. As a result of this, no new carbon is added to the natural carbon cycle [2].

Biomass includes only plant materials and recently dead biological species. It does not include organic materials that have been transformed by geological processes over many million years (such as petroleum or coal). Biomass can be obtained from agricultural, forest, municipal and biological products [2].

2.1 Constituents of Biomass

The constituents of biomass are associated with a variety of physical forms. For plant-derived biomass, the main structural cell constituents are cellulose, hemicellulose and lignin. The ratio of these three polymers varies from one plant to another. Other major bio-organic polymer structures such as starch and chitosan can also be found [3].

2.1.1 Cellulose

Cellulose is the most common organic polymer on Earth. It is a homopolysaccharide [3] (퐶6퐻10푂5)푛 of glucose 퐶6 sugar units. It is the main constituent of the cell walls of plants and comprises of about 33 percent of all vegetable matter [4]. Cellulose is indigestible by man but can be digested in herbivores with the help of micro-organisms present in their alimentary canal [4]. Cellulose is a major component in the raw materials of many manufactured goods such as fibers, paper etc. [3]. It can also be chemically modified to yield other substances which are useful in the manufacture of plastics, rayon, photographic films [4] etc.

Cellulose make up about 40-50wt% of most biomass materials on a dry fuel basis [3]. It is tasteless, odorless and insoluble in water and most organic solvents. At high temperatures, concentrated mineral acids can be used to break down cellulose chemically into glucose units [5].

Cellulose structure consists of up to 14,000 linearly coupled 퐷 − 푔푙푢푐표푝푦푟푎푛표푠𝑖푑푒 units which are connected by 훽 − 푔푙푦푐표푠𝑖푑𝑖푐 linkages in a 1:4 fashion (Mohan et al., 2006; Sjöström,1993; Solomons, 1984). The figure below shows the molecular structure of cellulose with repeating units consisting of two glucose units called “cellobiose”. It has a high molecular weight with a molecular mass typically of the order 106 푘푔.푘푚표푙−1. The 훽 − 푔푙푦푐표푠𝑖푑𝑖푐 linkages are commonly referred to as weak bonds. These bonds are easily broken and help to initiate the degradation of the cellulose molecule [3].

Fig 2.1: Molecular structure of cellulose [2]

Unlike starch, cellulose is a straight chain polymer with no coil or branches. It is also more crystalline than starch [5].

2.1.2 Hemicellulose

Hemicellulose is the second most abundant polymer. It contains about 20-50% of lignocellulose biomass. Hemicellulose is not chemically homogeneous like cellulose. It has branches with short lateral chains which consist of different types of sugars. These sugars (monosaccharides) include pentoses (rhamnose, xylose and arabinose), hexoses (glucose, galactose and mannose) and uronic acids (D-glucuronic, 4-O-methylglucuronic and D-galactouronic acids). The backbone of hemicellulose is either a hetero- or homo-polymer with short branches like beta-(1,4)-glycosidic bonds and occasionally beta-(1,3)-glycosidic bonds. Hemicellulose can have some degree of acetylation; applicable to heteroxylan. They have a lower molecular weight compared to cellulose. Hemicelluloses also differ in composition. Agricultural biomass such as straw and grasses are mainly composed of xylan while softwood hemicelluloses contain only glucomannan. [6].

Fig 2.2: Molecular structure of a typical hemicellulose, xylan [2]

2.1.3 Lignin

Lignin is the third most abundant polymer. It has a large and complex molecular structure containing cross-linked polymers of phenolic monomers. It is present in plants’ cell walls and improves structural strength, impermeability and resistance to microbial attack. Lignin serves as glue that binds the lignocellulosic biomass together making it impermeable to water. Herbaceous plants such as grasses usually have the least content of lignin while softwoods have the highest lignin content [6,7]. A typical hardwood contains 18 to 25% lignin, while a softwood contains 25 to 35% lignin by weight [2].

Fig 2.3: Some structural units of lignin [2]

2.2. Pyrolysis

Pyrolysis is the thermochemical decomposition of biomass at elevated temperatures (in the absence or little supply of oxygen) into a range of useful products. It involves the breaking down of large complex hydrocarbons into relatively simpler molecules of gas, liquid and char. It is one of several reaction steps observed in a gasifier. Pyrolysis of biomass usually takes place in a low temperature range of 300 to 650 C [2].

The composition of pyrolysis product depends on several factors most notably pyrolysis temperature and heating rate. Higher temperature and heating rate result in the formation of greater proportion of biogas and bio-oil while lower pyrolysis temperature and heating rate result in greater proportion of char. The initial pyrolysis product consists of condensable gases and solid char. The condensable gas can be further broken down into char, bio-oil and non- [2] condensable gases (CO, CO2, H2 and CH4 ) . The generic reaction of pyrolysis process is represented below:

ℎ푒푎푡 퐶푛퐻푚푂푝(퐵𝑖표푚푎푠푠) → ∑푙푖푞푢푖푑퐶푥 퐻푦푂푧 + ∑푔푎푠 퐶푎퐻푏푂푐 + 퐻2푂 + 퐶(푐ℎ푎푟) 1)

2.2.1. Pyrolysis Products

The pyrolysis of biomass produces three major products: liquid, solid and gas.

Liquid: The liquid is also called bio-oil. It is derived from the gases which have condensed to a liquid. It can be used as low grade diesel oil. It consists of tars, water and heavy hydrocarbons.

Solid: It is the solid material obtained from the pyrolysis process. It is also referred to as char. It consists of carbon. It can be used to improve soil fertility by preventing erosion and leaching.

Gas: These are the permanent (non-condensable) gases obtained from the pyrolysis process. It can also be referred to as syngas or biogas. They are several gases that make up the biogas. They include 퐶푂2, 퐻2푂, 퐶푂, 퐶2퐻2,퐶2퐻4, 퐶2퐻6,퐶6퐻6, 푒푡푐

2.2.2 Types of Pyrolysis

Pyrolysis can be broadly classified into slow, flash and fast pyrolysis based on the processing time and temperature of biomass [8].

Slow Pyrolysis Slow pyrolysis is characterized by low temperatures, slow biomass heating rates and lengthy solids and gas residence times. In slow pyrolysis, the heating temperature ranges from 0.1 to 2℃ (32.2 to 35.6℉) per second and the biomass is subjected to temperatures near 500℃ (932℉). The gas residence time may be over five seconds while that of biomass may range from minutes to days. During slow pyrolysis, char is released as the main product of the pyrolysis process [8].

Flash pyrolysis Flash pyrolysis occurs at moderate temperatures between 400 and 600℃ (752 and 1112℉) and very rapid heating rates. The vapor residence time in flash pyrolysis is less than 2 seconds. Flash pyrolysis produces fewer amounts of tar and gas when compared to the slow pyrolysis process [8].

Fast Pyrolysis In fast pyrolysis, the main products of the pyrolysis process are biogas and bio-oil. The biomass is heated to about 500℃ or higher depending on the desired amount of bio-oil or biogas. The peak temperature can be up to 1000℃ if biogas is the product of interest. Fast pyrolysis is characterized by i) High heating rate ii) Reaction temperature within 425 to 600℃ iii) Short Vapor residence time of <3seconds iv) Rapid quenching of the product gas [2]

2.2.3 Pyrolysis Product Yield

The product distribution of pyrolysis depends on several physical, chemical and operating parameters used in the pyrolysis process. Some of physical, chemical factors and operating parameters and their descriptions are outlined below a) Physical factors: Particle size b) Chemical factors: Biomass Composition c) Operating parameters: Heating rate and residence time, pyrolysis temperature

Particle Size The size of biomass particle has a significant influence on the yield of the pyrolysis process through their effect on the heating rate. When smaller sized particles are heated, they offer less resistance to the escape of condensable gases from the biomass to the surroundings before the gases undergo secondary cracking. This results in a higher liquid yield due to an increased escape of condensable gases. Larger sized particles on the other hand offer a greater resistance to the escape of condensable gases. This facilitates secondary cracking thus creating more biochar [2].

Biomass Composition The composition of biomass, especially its hydrogen-to-carbon (H/C) ratio, also has an influence in the product yield. The three major constituents of a lingo-cellulosic biomass have its specified temperature range of decomposition. Kumar et al., demonstrated that the thermogravimetric analysis (TGA) Differential thermogravimetry (DTG) data on some selected biomass shows the below range of temperatures for the initiation of pyrolysis [2] Hemicellulose: 150 to 350℃ Cellulose: 275 to 350℃ Lignin: 250 to 500℃ Each constituent undergo pyrolysis differently under different temperature ranges thereby producing different pyrolysis yields. For example, the main sources of volatiles in lingo- cellulosic biomass are cellulose and hemicellulose. Cellulose is the main source of condensable vapor. Lignin is a major contributor to char yield [2].

Heating rate and Residence time The heating rate of the biomass also has an effect on the products obtained from the biomass pyrolysis process. Rapid heating to moderate temperatures with range from 400 to 600℃ results in higher volatiles which increases liquid yield. Slower heating to same moderate temperature range of 400 to 600℃ yields more char [2]. Debdoubi et al. showed that increasing the heating rate increased the liquid yield [2]. Aside from the heating rate, the product residence time in the reactor also influences the yield. For slow pyrolysis, gradual removal of volatiles from the

16 reactor causes a secondary reaction between the volatiles and char leading to the formation of a secondary char. The heating rate and residence time of a pyrolyzer are adjusted to meet the desired product yield. The design norm for maximizing each of the pyrolysis products are given below [2]: i) A slow heating rate (<0.01 to 2.0℃/푠), a low final temperature and long gas residence time is used to maximize char production. ii) A high heating rate, a moderate final temperature (450 to 600℃) and a short gas residence time is used to maximize liquid production. iii) A slow heating rate, a high final temperature (700 to 900℃), and a long gas residence time maximize gas production.

Pyrolysis Temperature The maximum temperature the biomass particle is heated to and sustained is called the pyrolysis temperature. The pyrolysis temperature also has an effect on the composition and yield obtained from the biomass process. Low pyrolysis temperatures result in the formation of more char and less volatile while higher temperatures result in less char and more volatiles as applicable in gasification process.

Chapter 3: Types of Pyrolyzers for Biomass Pyrolysis

Pyrolyzers have historically been used to produce charcoal. Early pyrolyzers maximized charcoal production by employing slow heating rate, longer residence time and were operated in batch process mode. The pyrolysis temperature and duration of pyrolysis were also adjusted accordingly to obtain the desired product composition and yield [2].

Modern pyrolyzers may be operated in continuous or batch mode with emphasis on producing gaseous and liquid products. The type of reactor to be used in each pyrolysis application is dependent on the type of pyrolysis and the heat transfer requirement. There has been improvement in pyrolysis design, development and implementation over the last twenty five years. Based on the gas-solid contacting mode, pyrolyzers can be classified into fixed bed, fluidized bed and entrained bed. It can be further subdivided depending on the pyrolysis type the design is intended.

Fast pyrolysis reactors include

a) Fluidized Bed Reactors (Bubbling and Circulating) b) Rotating Cone Reactors c) Ablative Pyrolysis Reactors d) Twin Screw Reactors e) Ultra-Rapid Pyrolyzers

Slow Pyrolysis reactors include

a) Fixed Bed Pyrolyzers b) Fluidized Bed Reactors c) Vacuum Pyrolysis Reactors d) Heated Kiln Reactors e) Screw Feeders/Auger Reactors

3.1 Fast Pyrolysis Reactor a) Fluidized Bed Reactors

This can be divided into bubbling fluidized bed and circulating fluidized bed.

i) Bubbling Fluidized Bed Reactors

Bubbling fluidized bed reactors have been in operation for over fifty years. They have been used extensively in the petroleum and chemical processing industries. They are characterized by their provision of uniform and high heat transfer rate to the biomass [9]. Rapid particle heating is

18 achieved by pumping a hot gas (an inert gas or a flue gas) upward through the biomass particles by means of a porous distributor plate. When the gas velocity reaches the minimum fluidization velocity, the bed particles will behave like a fluid [10].

The required heat for the pyrolysis process is provided by burning a part of the gas in the bed (direct) or by burning the solid char in a separate chamber (indirect) and transferring the heat to the bed [2]. By selecting the appropriate size for the fluidized bed particle, the gas flow rate and the vapor residence time can be set to the desired level. Experience has shown that operating temperatures of 500- 550C and a vapor residence time of 0.5seconds resulted in the highest liquid yields [9]. For dry wood feed, the pyrolysis product typically contains about 70 to 75% liquid [2].

Fig 3.1: Bubbling fluidized bed pyrolyzer [2]

ii) Circulating Fluidized bed (CFB) Pyrolyzer

The circulating fluidized bed works on the same fluidization principle as the fluidized bed reactor except that the bed is highly expanded and the bed particles circulate in a constant loop comprising of a cyclone and loop seal [2]. Rapid and uniform heating of the bed particles around the entire height of the unit are distinguishing characteristics of this type of pyrolyzer. Several configurations of the circulating fluidized bed exist. But the different designs generally include a reactor and a cyclone. In CFB pyrolyzer, gas and solids move up the reactor with some degree of internal refluxing. The average residence time of solid particles is longer than the gas residence time though this difference is not as high as in a bubbling fluidized bed [2]. A major advantage of a CFB is that the entrained char from the reactor can be easily separated and burnt which can re- supply heat back to the reactor [2].

Fig 3.2: Circulating Fluidized Bed Reactor [2] b) Rotating cone reactors

The rotating cone pyrolyzer reactor differs from the fluidized bed reactor in that a centrifugal force is used to transport the biomass particles and products in a spiral through the reactor in either an upward or downward direction. Hot carrier solids mixed with biomass transfers the heat and are separated from the products after pyrolysis and recycled as shown by Wagenaar et el. [10]. The biomass is fed to the bottom of the rotating cone at about 360 to 960 rev/min. The biomass undergoes rapid heating of about 5000K/s and is pyrolyzed within a small annular volume. The condensable biogas leaves through another tube, while the sand and char spill over the upper rim of the cone into the fluidized bed surrounding it. The char is burnt in the fluidized bed and the heat produced is used to heat the cone and the recycled solid to supply the heat for pyrolysis. The distinguishing features of this pyrolyzer are very short solids and gas-phase residence time of about 0.5 seconds and 0.3 seconds respectively. Typical liquid yield is 60 to 70% as shown by Hulet et al[2].

Fig 3.3: Rotating cone pyrolyzer reactor [2]. c) Ablative Pyrolyzer

The ablative pyrolysis reactor involves pressing the biomass against a hot moving surface either by a centrifugal force or mechanical means (push/press). This creates a high pressure between the biomass particle and the hot reactor wall which ensures rapid heat transfer to the biomass particle. The liquid melts out of the biomass and the final liquid yield may be as high as 80% as shown by Diebold et al. This pyrolyzer type possesses high heat transfer and short gas residence time [2]. The advantage of this pyrolyzer type is that it does not require a carrier gas.

Fig 3.4: Ablative pyrolyzer [2]

21 d) Twin Screw Reactors

The twin screw pyrolysis reactors involve the use of two offset screws encased in the reactor. These screws fluidize and transport the biomass particles in the reactor. They rotate in the same direction and have intertwining flights to dislodge particles caught in the thread as shown by Raffelt et al. The biomass is mixed with a heat carrier such as sand at the reactor entry which provides heat for pyrolysis. The gas or char produced may be combusted to reheat the heat carrier material. Rafflelt et al. showed experimentally that the product obtained from this pyrolysis type contains 53 to 78% liquid, 8 to 20% gas and 12 to 34% char.

Fig 3.5: Twin screw pyrolysis process [10]

f) Ultra-Rapid Pyrolyzers

Ultra-rapid pyrolyzer is characterized by high heating rate and short residence time for producing high liquid yield. Inert gas (such as nitrogen) is heated to 100℃ above the temperature of the reactor at a very high velocity and sent into the reactor. The nitrogen bombards a stream of biomass injected into the reactor. A heat carrier (such as sand) can also be heated externally and used to bombard a stream of biomass through multiple jets. This creates an exceptionally high heating rate which heats the biomass to the pyrolysis temperature in a few milliseconds. The products obtained from pyrolysis leaves the reactor from the bottom and is immediately cooled off to prevent secondary reactions. This process therefore maximizes liquid yield [2].

Fig 3.6: Ultra-rapid pyrolyzer [2]

3.2 Slow pyrolysis reactor

These reactors usually operate under slow heating rate and lower biomass pyrolysis temperature. Some of these reactors are described below:

a) Fixed Bed Pyrolyzers

This is the oldest type of pyrolyzer. It typically works in batch mode. The heat for pyrolysis is supplied by an externally source or through limited combustion in a beehive oven. Due to volume expansion, the product may flow out of the pyrolyzer while the char remains in the pyrolysis reactor. A sweep gas can be used for effective removal of the product gas from the pyrolysis reactor. This pyrolyzer typically operates with slow heating rate and long residence time of the product. Char is the main component obtained from this type of pyrolysis [2].

Fig 3.7: Beehive oven as a fixed pyrolyzer for charcoal production [2]

b) Fluidized Bed Reactors

Similar in operation to Fluidized bed for fast pyrolysis but has a longer residence time, lower pyrolysis temperature and allows for a greater proportion of charcoal production. It can generate product ratios of around 35 to 45% liquid, 30 to 35% gas and 20 to 25% charcoal as shown by Zabanitou et al. [10].

c) Vacuum Pyrolysis Reactors

Vacuum pyrolysis is a slow pyrolysis process with a lower heating rate and residence time of approximately 40s. Vacuum pyrolysis utilizes low absolute pressure inside the reactor. The biomass is fed into the reactor by a conveyor where it is mechanically stirred during the pyrolysis process. The product (charcoal) is expelled from the reactor through a pressure seal. The liquid yield is high due to the low pressure conditions reducing the reaction rates and condensation reactions in the vapor phase as shown by Effendi et. al. [10].

d) Heated Kiln Reactors

The heated kiln reactor is characterized by much longer residence time compared to vacuum pyrolysis reactors and fluidized bed reactors. They are the most common reactor design for slow

24 pyrolysis. They may be heated externally by a heating jacket surrounding the kiln or internally by the combustion of product gases as shown by Honnery et al. It may be operated in batch mode which does not require agitation or operated in continuous mode and require some type of mechanical transport. Several methods can be employed for mixing and transporting in the heated kilns such as conveyors or sweepers or the use of rotating or inclined kiln. All types offer a simple cost-efficient design. The product output may differ depending on several factors like the residence time and temperature [10].

Fig 3.8: Heated Kiln pyrolysis process [10]

e) Screw Feeders/Auger Reactors

Screw or auger reactors essentially work in the same principle as the heated kiln reactors. They consist of a rotating helical screw encased in a tube. The screw mixes and transports the biomass through the tube in a rotational motion. The advantage of screw feeder (auger) reactors is the efficient mixing and transport of the biomass particles through the reactor. The process conditions (such as residence time and reactor temperature) determine the product output. Liquid yield can range from 18 to 25 % and charcoal yields can range from 50 to 60 % as shown by Bridgewater et al. Day et al. showed particle size less than 1mm is required for ideal operation which means considerable cost for feed grinding is required. Blockage of the internals by biomass compaction is a disadvantage of screw feeder pyrolyzer [10].

Fig 3.9: Screw/Auger pyrolysis process [10]

Chapter 4: Fluidization and Design of a Fluidized Bed Reactor

Fluidized beds have gained wide usage in the petroleum and chemical processing industries. They are used for separations, rapid heat and mass transfer operations, and catalytic reactions. A typical fluidized bed is cylindrical in shape and contains particles through which fluid (gaseous or liquid) flows. For a fluidized bed reactor, sand can be used as a catalyst. The velocity of the fluid released to the reactor should be sufficiently high to suspend or fluidize the particles thereby providing a large surface area for the reaction. Fluidized beds can range in sizes from small scale laboratory equipments to large industrial systems [11]. Fluidized bed reactor for biomass pyrolysis was chosen for this research work because of the following advantages.

a) It involves a simple design b) The cost of design and construction is relatively cheap c) Easily repaired when faulty d) Efficient mixing of particles in the reactor e) Uniform temperature of particles in the reactor

Several fluidized bed reactor have been developed over the years for commercial purposes. Existing designs for commercial purposes include:

a) Fortum/Valmet fast pyrolysis plant in Finland b) AE Cote-Nord Bioernegy/Ensyn fast pyrolysis plant in Canada c) Dynamotive plant in Canada

A major limitation of existing designs is lack of mobility. The existing designs involve the construction of the reactor over a large area of land. The biomass needs to be transported to these locations before they can be processed. This makes the process complex, time consuming, expensive and requires adequate planning and logistics. The mobile fluidized bed reactor proposed in this research solves many of these issues. It can be easily moved from one place to another depending on locational need.

4.1 Fluidization theory

Fluidization occurs when an upward-flowing gas passing through a porous distributor plate, imposes a sufficiently high drag force to overcome the gravitational force exerted on the particles. The drag force is as a result of the frictional force imposed on the particle by the gas; the particle imposes an equal and opposite drag force on the fluidizing gas. This implies that as a particle becomes more fluidized, the local gas velocity around the particle due to these drag forces are also affected. For spherical particles, this effect is minimal, however, the effect of these drag forces is significant for irregularly shaped particles [12].

When the particles are fluidized, the bed behaves differently as velocity, solid and gas properties are varied. There are a number of flow regimes that can result depending on the velocity of the fluidizing gas. When the flow of the fluidizing gas passing through a bed of particles is increased continually, some particles may vibrate, but the height of the bed remains the same. This is called a fixed bed (Fig. 3.10A). Increasing the gas velocity further, a point is reached when the drag force by the upward flowing gas equals the weight of the particles. The bed voidage also increases slightly. This is the onset of fluidization and it is called the minimum fluidization (Fig

3.10B). The velocity at this point is called the minimum fluidizing velocity 푢푚푓. Increasing the velocity of the gas beyond the minimum fluidizing velocity will lead to the formation of bubbles (Fig 3.10C). If the velocity is increased further, the bubbles coalesce and grow in size as they rise. If the ratio of the bed height to the bed diameter is sufficient, the diameter of the bubbles may be the same as the diameter of the bed. This is known as slugging (Fig 3.10D). At sufficiently high gas flow rate, the velocity exceeds the terminal velocity (the velocity required to transport particles out of the bed). The upper surface of the bed disappears and a turbulent motion (Fig 3.10E) of solid clusters and voids of gas of various shapes and sizes form in the bed. Further increase of the gas velocity creates an entrained bed with disperse, lean or dilute phase fluidized bed which results in pneumatic solid transport (Fig 3.10F) [13].

Fig 4.1: Schematic representation of the different regimes of fluidized bed by Kunni and Levenspiel

4.2 Geldart Classification of Particles

Geldart (1972) described the properties of particles and their ease of fluidization. This description is important for predicting the behavior of a powder (particle) when fluidized. According to Geldart, particles can be classified into four groups: A, B, C and D [14]. i) Group A

Materials are aeratable. They have small mean particle size (푑푝 < 30휇푚) and/or low particle 1.4푔 density(< ~ ). Fluid cracking catalysts belong to this category. Solid materials are easily 푐푚3 fluidized. Smooth fluidization occurs at low gas velocities without bubble formation. Bubbles are formed at higher gas velocity. The minimum bubbling velocity is always greater than the [13] minimum fluidization velocity(푢푚푏 > 푢푚푓 ) .

ii) Group B

Particles are ‘sandlike’. Particle size in this group range from 150 푡표 500휇푚 and density from 1.4 to 4 푔/푐푚3. Bubbles form when the gas velocity exceeds the minimum fluidization velocity. The bubbles formed can grow to a large size. Coarse sand and glass beads are common materials that belong to this group [13]. iii) Group C

Materials in this group are cohesive or very fine particles. Their size are usually less than 30휇푚. Because the interparticle forces are relatively large, compared to those resulting from gas action, they are extremely difficult to fluidize. Bubble formation does not occur. Talc, flour and starch are examples of materials that belong to this group [13]. iv) Group D

Materials in this group are called spoutable. The materials are either very large or very dense. They are undesirable for fluidization. As velocity of gas increases, a jet may be formed in the bed and bed materials can be blown out in a spouting motion with the jet. Spouting behavior and severe channeling can occur if the gas distribution is uneven. Lead shot, roasting coffee beans and some roasting metal ores are examples of group D materials [13].

Geldart’s classification offers a clean and easy way to classify particles for fluidization operations. For any particle of known density 휌푠 and mean particle size 푑푝, the fluidization characteristics can be predicted from the graph below. Other properties such as the bubble size, bubble velocity and existence of slugs, can also be predicted [13].

Fig. 4.2: Diagram of Geldart particle classification

4.3 Design of a Fluidized Bed Reactor

4.3.1 Material Selection

Table 4.1: Material properties and justification Poisson Young Part of reactor Material Justification ratio Modulus  Light material Handle Natural Rubber ~0.49 [15] 5MPa [16]  Good thermal and electrical Insulator

 Non corrosive Reactor and Stainless Steel ~197GPa 0.24 [16]  Easy weldability Flange* AISI 304L [15]

Bolt**, Nut***  Good ductility and Medium Carbon and 0.29 [18] 210GPa [17] strength Steel washer****  Wear resistant  Excellent chemical resistance against Chlorosulphonated Gasket 0.46[20] 15MPa[21] acids, alkalis and oil polyethylene  Good fire protection properties [19]

* Flange - 22 inches Nominal pipe size, 300lb Lap joint and Stub End Flange; **Bolt - Grade 8 bolt, size 1-1/2’’ (ASME B18.2.1-1996); ***Nut – size 1-1/2’’ (SAE); ****Washer SAE – size 1-1/2’’ (SAE)

Table 4.2 Reactor parameters Parameters Value Justification Design Pressure, 푃 0.92푀푃푎 = 9.2푏푎푟 = Arbitrarily chosen (퐴푠푠푢푚푒푑) 133.4 푝푠𝑖 as the Initial design value

Maximum Design Temperature, T 1500 Arbitrarily Chosen Allowable (°C) as a safe temperature. Temperature ideally should not exceed 1000°C Factor of Safety, 푓푠 4 To be able to withstand 4 times its design load for safety reasons Diameter of reactor, 퐷 0.5푚 Arbitrarily chosen Other Parameters (퐴푠푠푢푚푒푑) as the Initial design value Joint Efficiency, E 0.60 [22] Single butt weld with no backing strip

Reactor Wall thickness

Two main stresses occur on the shell portion of the fluidized bed reactor: the hoop stress and the longitudinal stress. The wall thickness necessary for hoop stress and longitudinal stress will be calculated. The greater of the two thicknesses will be chosen as the wall thickness.

For Hoop Stress,

푃×퐷 Wall thickness, 푡 = 2) 푤 2푆×퐸−1.2푃

For Longitudinal stress,

푃×퐷 Wall thickness, 푡 = 3) 푤 4푆×퐸+0.8푃 Where P is the design pressure (maximum allowable)

D is the reactor diameter

S is the maximum allowable stress. It is also denoted as 휎푚.

E is the joint efficiency

1 ′′ The minimum wall thickness (excluding corrosion allowance) is = 1.5875푚푚 [23]. 16 Corrosion allowance is always added to the wall thickness value incase corrosion of material occurs. The corrosion allowance for the reactor typically ranges from 1 ′′ 3 ′′ (1.5875푚푚) 푡표 (4.7625푚푚). 16 16

Maximum allowable stress (푆 표푟 휎푚) ,

푌푖푒푙푑 푆푡푟푒푛푔푡ℎ 표푓 Stainless Steel AISI 304L 푆 = 휎 = 푚 퐹푎푐푡표푟 표푓 푠푎푓푒푡푦 4)

210 × 106 푁/푚2 106푁 = = 52.5 × 4 푚2 Substituting the various parameter values into the hoop and longitudinal stress wall thickness gives

Hoop Stress,

푃퐷 106×0.5 Wall thickness, 푡 = = = 0.0081푚 = 8.1푚푚 푤 2푆퐸−1.2푃 2×52.5×106×0.6−1.2×106 Longitudinal Stress,

푃퐷 106×0.5 Wall thickness, 푡 = = = 0.0039푚 = 3.9푚푚 푤 4푆퐸+0.8푃 4×52.5×106×0.6+0.8×106

∴The minimum wall thickness 푡푤 = (8.1 + 4.7625)푚푚 = 12.8625푚푚 ≈ 13푚푚

4.3.2 Design of Distributor plate

The distributor plate is the base that holds the contents of the reactor (biomass and sand). It is a porous material allowing for easy gas flow through it. It divides the reactor into pyrolysis chamber above and the plenum below. The selection and design of the distributor plate is one important factor that determines the success of the fluidized bed reactor. A good distributor plate design ensures a stable fluidization and a uniform distribution of fluidizing gas to the reactor. Particle flow back during reactor shut down and start up should also be prevented by the distributor plate. The distributor plate remotely controls the hydrodynamics of the solid-gas flow, the mixing quality and size of bubbles. The distributor plate material should be carefully selected to meet the thermal and structural requirements of the reactor. It should be able to withstand high temperatures from the plenum (windbox) below and not susceptible to buckling due to the weight above.

Different types of distributors have been explored for different uses by designers. Common types of air distributors are [24] a) Nozzle or bubble cap distributor b) Porous and straight hole orifice distributor c) Sparge type distributor

Fig 4.3: Different types of distributors [24]

For successful operation of the fluidized bed reactor, the following are the desired properties of the gas distributor [24]

1. Stable and Uniform fluidization at all times during operation 2. Minimum dead zone areas on the distributor 3. Zero to negligible flow back of solid to the plenum chamber below 4. Minimum erosion of heat exchanger tubes and nearby nozzles in the bed

For optimum operation of the fluidized bed reactor, the distributor plate should be carefully selected and designed to achieve stable and uniform fluidization and prevent particle flow back during shut down. The methodology for designing the distributor plate and its nozzles are outlined below. The following parameters should be calculated in the order shown below:

i) Archimedes number, 퐴푟 The Archimedes number is used to determine the motion of fluids due to differences in densities. It is dimensionless and defined as the ratio of external forces to internal viscous forces [25]. The Archimedes number can be calculated by 3 휌푔푑푝(휌푝−휌푔)푔 퐴푟 = 5) 휇2

The Archimedes number is useful in estimating the minimum fluidization velocity, 푢푚푓 for the fluidized bed reactor.

ii) Reynolds number at the point of minimum fluidization, 푅푒푝,푚푓 Wen and Yu [26] proposed a second-order polynomial that can be used to calculate the Reynolds number at the point of minimum fluidization velocity.

2 퐴푟 = 1,650푅푒푝,푚푓 + 24.5푅푒푝,푚푓 6)

iii) Minimum Fluidization Velocity, 푢푚푓 The minimum fluidizing velocity can be calculated from the Reynolds number obtained at the point of minimum fluidization 휇푅푒푝,푚푓 ∴ 푢푚푓 = 7) 휌푔푑푝

iv) Superficial velocity, 푢; Terminal velocity, 푢푡; Minimum Slugging velocity, 푢푚푠; voidage at minimum fluidization, 휀푚푓 and Bed height at minimum fluidization, 퐻푚푓

The above parameters are interrelated and therefore have to be determined simultaneously without compromising the design requirements.

The superficial velocity, 푢, is the velocity of the fluidizing gas entering the reactor. For proper operation of the fluidized bed reactor, the superficial velocity must be above the minimum [27] fluidizing velocity, but below the terminal 푢푡 and minimum slugging velocities 푢푚푠.

푢푚푓 < 푢 < 푢푡 푎푛푑 8)

푢푚푓 < 푢 < 푢푚푠 9)

Terminal velocity, 푢푡 is the velocity required to transport the particles out of the fluidized bed reactor. The formula for calculating the terminal velocity depends on the Reynolds number. 푔(휌 −휌 )푑2 푢 = 푝 푔 푝 푓표푟 푅푒 < 0.4 10) 푡 18휇

2 1/3 1.78×10−2×[푔(휌 −휌 )푔] 푝 푔 [3] 푢푡 = ( ) 푑푝 푓표푟 0.4 < 푅푒 < 500 11) 휌푔휇

Slugging velocity, Slugging occurs when the bubbles formed in the fluidized bed equals the bed/reactor diameter which results in the formation of short discrete cylinders. The minimum velocity for the onset of slug formation is called the minimum slugging velocity 푢푚푠. According to Geldart, the minimum slugging velocity can be obtained from the equations below [28]:

푢 + 0.07(푔퐷) 0.5 + 0.16(1.3퐷0.175 − 퐻 )2 퐻 < 1.3퐷0.175 { 푚푓 푚푓 푚푓 푢푚푠 0.5 0.175 12) 푢푚푓 + 0.07(푔퐷) 퐻푚푓 ≥ 1.3퐷

Voidage at minimum fluidization, 휀푚푓 This is the fraction of gas (space) available inside the bed particles at the point of minimum fluidization. This can be calculated by using the Todes’ equation [29]:

휀 휀푚푓 = 2 0.21 13) 푅푒푝+0.02푅푒푝 [ 2 ] 푅푒푝,푚푓+0.02푅푒푝,푚푓

Bed Height at minimum fluidization, 퐻푚푓 This is the height of the bed particles at the point of minimum fluidization. In order to calculate this parameter, we need to utilize the Davidson two-phase model of the fluidized bed which implies that all particles in the fluidized bed are outside the bubbles formed and the bed is always in a state of minimum fluidization. The rationale behind this assumption is that the porosity of the particulate phase is equal voidage at minimum fluidization, 휀푚푓, and therefore volume fraction occupied by bubbles is solely responsible for any increase in bed height [29]. Therefore,

ℎ푝(1−휀) 퐻푚푓 = 14) 1−휀푚푓

Where ℎ푝 is the height of particles in the reactor

v) Determination of the Bed Pressure drop ∆푃퐵 The bed pressure drop across the fluidized bed is highest at the point of minimum fluidization. Increase in superficial gas velocity beyond this point does not increase the bed pressure drop. In most cases, the value of the fluidizing gas velocity (superficial velocity) 푢 is higher than the minimum fluidizing velocity 푢푚푓 but lower than the terminal velocity, 푢푡. For a packed bed, the pressure drop across the bed can be calculated from the Ergun Equation. For a fluidized bed, the pressure drop across the bed (beyond the point of minimum fluidization), is calculated from the equation below:

∆푃퐵 = ℎ푝(1 − 휀)(휌푝 − 휌푔 )푔 15)

vi) Distributor plate pressure drop, ∆푃퐷 The pressure drop across the distributor plate is very important consideration in the design of a fluidized bed reactor. A suitable pressure drop ensures uniform and stable fluidization. Several [24] correlations for distributor plate pressure drop exist. Kunii et al. proposed ∆푃퐷 ≈ (0.15 − 0.40)∆푃퐵. Most designers however use a distributor plate pressure drop that is 30% the bed pressure drop i.e.

∆푃퐷 = 0.3∆푃퐵. 16)

vii) Gas velocity through an orifice, 푢표

The gas velocity through an orifice is the first parameter that needs to be calculated to develop a correlation between the number of holes on all the orifices and the pressure drop across the distributor plate. The gas velocity through an orifice can be calculated from [24]

0.5 2∆푃퐷 푢표 = 퐶퐷 [ ] 17) 휌푔 푎푡 표푟푖푓푖푐푒

where 퐶퐷 is the orifice coefficient. Zenz et al. suggested the orifice coefficient to be 0.8. For a thick plate 푡 푡 0.13 > 0.09, the correlation of Quereshi and Creasy (1979) where 퐶퐷 = 0.82( ) may be used. From 푑표푟 푑표푟 erosion standpoint, orifice velocities that are less than 30m/s are considered to be safe as shown by Pell et al., while orifice velocities exceeding 90m/s are considered to be risky as shown by Geldart et. al. [24]

viii) Number of orifices per unit area, 푁, number of orifices per nozzle, 푛표, number of nozzles, 푛, diameter of nozzle, 푑푛, and pitch, 푃

푁 is calculated from the formula for the fractional opening of the orifice.

푢휌 [24] 푔 4 Fractional opening of the orifice 푁 = 2 18) 푢표휌푔 푎푡 표푟푖푓푖푐푒 휋푑표

With 푑표 determined (fixed) by the designer, N can be calculated from the formula above.

As a rule, for the design of the nozzle, the area of the internal diameter of each nozzle must be greater than the sum of the areas of the orifices of that nozzle [24].

휋푑2 휋푑2 푛 표 < 푛 19) 표 4 4

Where 푑푛 is the internal diameter of the nozzle.

푁 The number of orifices per nozzle is calculated from 푛 = 20) 표 푛

The Pitch, 푃 is the distance between the midpoints of two adjacent nozzles.

2 For triangular pitch, 푁 = 21) √3푃2

1 For square pitch, 푁 = 22) 푃2

ix) Location of bed internals relative to jet length/height

Erosion of bed internals is major consideration in the design of a fluidized bed. Bed particles carried by the gas collide with the walls of the reactor thereby causing erosion. Nozzles should be located at a safe distance to minimize erosion of the walls. The jet configuration can be vertical, horizontal or directed downward as shown in the figure below [30]. Horizontally directed jets prevent particle flow back during shut down therefore they will be used in this research work.

Fig 4.4: Nozzle Jet configurations [30]

The approximations of the relation for the different configuration is given by Karri et al as [31, 32].

퐿푢푝~2퐿ℎ표푟 ~3퐿푑표푤푛 23)

Merry (1975) proposed a correlation for determining the jet height 퐿푢푝. To minimize erosion of bed internals due to the high velocity of the gas entering the reactor, the correlation for jet height [24] 퐿푢푝 below is considered safe in the design of the reactor .

0.3 0.2 퐿 휌 푑 푢2 푢푝 = 5.2 [ 푔 표] [1.3 ( 표 ) − 1] 24) 푑표 휌푝푑푝 푔푑표

The above correlation when used in combination with the correlation by Karri (1991) will give an approximate value for the various configurations.

x) Minimum Nozzle height 푙푚

To minimize attrition of bed internals, the gas entering the nozzle should have sufficient room for expansion. Increasing the nozzle height, is a very effective way to achieve this requirement. [24] According to Zenz et al., the minimum nozzle length, 푙푚 is given as

푑 −푑 푙 (𝑖푛 푚푒푡푒푟푠) = 푛 표 25) 푚 0.193

xi) Thickness of the distributor plate, 푡푝

The plate must be strong enough to support the weight of the biomass particles without buckling. The distributor plate will be regarded as a circular plate subjected to uniform loading with edges clamped to the wall of the reactor. To obtain the thickness of the plate, the maximum deflection for the plate (at the center) will be assumed and used to calculate the plate thickness.

For circular plate under uniform loading with the edges simply supported, the maximum stress at the center is

3(3+푣)푃푟2 휎푚 = 2 26a) 8푡푝

The thickness of the plate required can be determined from the above equation where all other variables are known.

2 3(3+푣)푃푝푙푎푡푒 푟 푡푝 = √ 26b) 8휎푚

xii) Distributor plate nozzle thickness The minimum thickness for pressure vessels (walls) excluding corrosion allowance is 1 ′′ = 1.5875푚푚 [10]. This value will be used for the minimum nozzle thickness. 16 The above values for the steps explained in i) to xii) is calculated below:

Table 4.3 Reactor Design Analysis Parameters Parameters Value Diameter of reactor, 퐷 (퐴푠푠푢푚푒푑) 0.5푚 Initial Height of sand in the reactor, ℎ푠 0.1푚 Initial Height of biomass in the reactor, ℎ푏 0.7푚 Initial Height of particles in the reactor, ℎ푝 0.8푚 3[1] Density of biogas, 휌푔 1.15푘푔 /푚 3 [2] Density of sand, 휌푠 1281 − 1602푘푔/푚 * Density of biomass (dry wood chips), 0.38푔 = 380푘푔/푚3[3] 푐푚3 3 Density of particles (sand and biomass), 휌푝 532.72푘푔/푚 (Calculated below) Diameter of particle, 푑푝 (퐴푠푠푢푚푒푑) 0.2푚푚 Acceleration due to gravity, 푔 9.81푘푔/푚3 Dynamic viscosity, 휇 2.11 × 10−5 푘푔/푚푠**[4] Natural porosity of fine sand, 휀 0.29 − 0.46[5]***; ퟎ.ퟓ 풘풊풍풍 풃풆 풖풔풆풅 3 *The maximum density of sand (1602 푘푔 ) will be used to account for the maximum possible weight 푚

** Due to the difficulty of obtaining the actual dynamic viscosity of biogas at 400℃, the dynamic viscosity of methane at 400℃ was used since it the major constituent of biogas (푎푏표푢푡 50− 75%)[2]. ***Porosity of 0.5 will be used for the individual particles (sand and biomass)

Density of particles (sand and biomass) Since the reactor contains sand and biomass, the above approach will be taken to determine the combined densities of particles (sand and biomass), 휌푝 in the reactor.

푡표푡푎푙 푚푎푠푠 푚푎푠푠 표푓 푠푎푛푑+푚푎푠푠 표푓 푏푖표푚푎푠푠 휌 = = 27) 푝 푡표푡푎푙 푣표푙푢푚푒 푐표푚푏푖푛푒푑 푣표푙푢푚푒

푚푎푠푠 표푓 푠푎푛푑, 푚푠 = 푑푒푛푠𝑖푡푦 표푓 푠푎푛푑, 휌푠 × 푣표푙푢푚푒 표푓 푠푎푛푑 𝑖푛 푟푒푎푐푡표푟,푣푠 28)

퐷2 = 휌 푣 = 휌 휋 ℎ (1 − 휀) 29) 푠 푠 푠 4 푠 0.52 = 1602 × 휋 × × 0.1 × (1 − 0.5) 4

= 15.73푘푔

푚푎푠푠 표푓 푏𝑖표푚푎푠푠, 푚푏 = 푑푒푛푠𝑖푡푦 표푓 푏𝑖표푚푎푠푠,휌푏 × 푣표푙푢푚푒 표푓 푏𝑖표푚푎푠푠 𝑖푛 푟푒푎푐푡표푟, 푣푏 30)

퐷2 = 휌 푣 = 휌 휋 ℎ (1 − 휀) 31) 푏 푏 푏 4 푏 0.52 = 380 × 휋 × × 0.7 × (1 − 0.5) 4

= 26.11푘푔

푡표푡푎푙 푚푎푠푠 표푓 푝푎푟푡𝑖푐푙푒푠 = 푚 = 15.73 + 26.11 = 41.84푘푔 32)

푡표푡푎푙 푚푎푠푠 15.73 + 26.11 41.84 휌 = = = = 532.72푘푔/푚3 푝 푡표푡푎푙 푣표푙푢푚푒 퐷2 0.52 휋 ℎ (1 − 휀) 휋 × × 0.8 × (1 − 0.5) 4 푝 4

i) Determination of the Archimedes number

From equation 5,

1.15 × 0.23 × (532.74 − 1.15) × 9.81 퐴푟 = 10003 × (2.11 × 10−5푘푔/푚푠)2

퐴푟 = 107.76 ii) Reynolds number at the point of minimum fluidization determination From Wen and Yu [26] in equation 6,

퐴푟 = 107.76

Solving the above polynomial to find 푅푒푝,푚푓 gives

푅푒푝,푚푓 = 0.06525 iii) Calculation of Minimum fluidization velocity, terminal and slugging velocities

Recall from equation 7,

2.11 × 10−5푘푔/푚푠 × 0.06525 ∴ 푢 = 푚푓 0.2 1.15푘푔/푚3 × 푚 1000

−1 푢푚푓 = 0.006푚푠 iv) Determination of the Superficial velocity, 푢; Terminal velocity, 푢푡; Minimum Slugging velocity, 푢푚푠; voidage at minimum fluidization, 휀푚푓 and Bed height at minimum fluidization, 퐻푚푓

As stated previously, the above parameters are interrelated and therefore have to be determined simultaneously without compromising the design requirements.

The value for the superficial gas velocity, 푢 will be conveniently chosen between the

values of minimum fluidization velocity,푢푚푓, and the terminal (푢푡) and minimum

slugging velocities (푢푚푠) i.e. the superficial velocity (푢) must be greater than the minimum fluidizing velocity 푢푚푓 but less than the terminal (푢푡) and slugging velocities (푢푚푠).

Recall that

−1 푢푚푓 = 0.006푚푠

Next we need to know the terminal and slugging velocities. To determine the terminal velocity, knowledge of the Reynolds number is required. −1 Guessing the superficial gas velocity 푢 greater than 푢푚푓 , 푠푎푦 푢 = 0.02 푚푠 helps us to determine other parameter. If this value of superficial velocity suits the all the design criteria, then it will be used. Solving for the Reynolds number, 휌푔푢푑푝 푅푒 = 33) 푝 휇

1.15 × 0.02 × 0.2 = = 0.218 2.11 × 10−5 × 1000

Since 푅푒푝 = 0.218 < 0.4, 2 푔(휌푝 − 휌푔 )푑푝 푢 = 푓푟표푚 푒푞푢.10) 푡 18휇 9.81 × (532.72 − 1.15) × 0.22 푢 = = 0.549 푡 18 × 2.11 × 10−5 × 10002

−1 Since the superficial gas velocity 푢 = 0.02푚푠 < 푢푡 = 0.549, it satisfies this criterion.

Next is the minimum slugging velocity, 푢푚푠 Recall from equation 12) [28] 푢 + 0.07(푔퐷) 0.5 + 0.16(1.3퐷0.175 − 퐻 )2 퐻 < 1.3퐷0.175 { 푚푓 푚푓 푚푓 푢푚푠 0.5 0.175 푢푚푓 + 0.07(푔퐷) 퐻푚푓 ≥ 1.3퐷 Since the minimum slugging velocity requires the value of the bed height at minimum fluidization 퐻푚푓 which has not been calculated yet, it will be determined at a later stage.

Voidage at minimum fluidization, 휀푚푓 [29] 휀푚푓 can be obtained from equation 13) 휀 휀푚푓 = 2 0.21 푅푒푝 + 0.02푅푒푝 [ 2 ] 푅푒푝,푚푓 + 0.02푅푒푝,푚푓

0.5 = = 0.388 0.218 + 0.02 × 0.2182 0.21 [ ] 0.06525 + 0.02 × 0.06525 2

Since the 휀 and 휀푚푓 are known, the bed height at minimum fluidization can be determined. From Davidson two-phase model assumption, Bed Height at minimum fluidization [29] 퐻푚푓 can be obtained from equation 14 , 0.8(1 − 0.5) 퐻 = = 0.654푚 푚푓 (1 − 0.388)

[28] Since 퐻푚푓 = 0.654 < 1.15, from equation 12 , 0.5 0.175 2 푢푚푠 = 푢푚푓 + 0.07(푔퐷) + 0.16(1.3퐷 − 퐻푚푓) = 0.006 + 0.07(9.81 × 0.5) 0.5 + 0.16(1.3 × 0.50.175 − 0.654)2 = 0.201

푢푚푠 = 0.201푚/푠

From the above calculations, it can be observed from equation 8 and 9 [27] that

푢푚푓 < 푢 < 푢푡 𝑖. 푒.0.006 < 0.02 < 0.549 푎푛푑

푢푚푓 < 푢 < 푢푚푠 𝑖. 푒.0.006 < 0.02 < 0.201 The value of 푢 = 0.02푚푠−1satisfies both criteria therefore it will be used in this design.

v) Determination of the bed pressure drop, ∆푃퐵 From equation 15,

∆푃퐵 = ℎ푝(1 − 휀)(휌푝 − 휌푔 )푔 = 0.8 × (1 − 0.5)(532 .72 − 1.15) × 9.81 = 2,085.88푁/푚2

vi) Distributor plate pressure drop, ∆푃퐷 From equation 16, ∆푃퐷 = 0.3∆푃퐵 = 0.3 × 2,085.88 = 625.76 푁/푚2

vii) Gas velocity through an orifice, 푢표 From equation 17 [24], 0.5 2∆푃퐷 푢표 = 퐶퐷 [ ] 휌푔 푎푡 표푟푖푓푖푐푒 2 × 625.76 0.5 푢 = 0.8 × [ ] = 26.39푚/푠 (푠푎푓푒 푓표푟 푑푒푠𝑖푔푛) 표 1.15

44 viii) Number of orifices per unit area, 푁, number of orifices per nozzle, 푛표, number of [24] nozzles, 푛, diameter of nozzle, 푑푛, and pitch, 푃. From equation 18 푢휌푔 4 푁 = 2 푢표휌푔 푎푡 표푟푖푓푖푐푒 휋푑표

The diameter of the orifice 푑표will be assumed to be 3푚푚 = 0.003푚 (chosen from reasonable estimate)

0.02 × 1.15 × 4 푁 = = 107.21 26.39 × 1.15 × 휋 × 0.0032 108표푟𝑖푓𝑖푐푒푠 푁 ≈ 푚2 Recall from equation 21 [24] that the relationship between the number of orifices per m2, N and the pitch (triangular)*, P is

2 푁 = √3푃2

2 푃 = √ = 0.1034푚 = 103.4푚푚 푁√3

∗More nozzles will be contained in the distributor plate for a triangular pitch than for a square pitch.

For a triangular pitch of 103.4 mm and a distributor plate area of ≈ 0.2푚2, there are 19 nozzles (obtained from manual counting from CAD drawing).

∴ 푛푢푚푏푒푟 표푓 푛표푧푧푙푒푠,푛 = 19

[24] The number of orifices per nozzle, 푛표, from equation 20 푁 108 푛 = = = 5.68 ≈ 6 표 푛 19

Diameter of nozzle, 푑푛

Recall that the area of the nozzle must be greater than the sum of areas of all the orifices on the nozzle [24]. From equation 19 [24],

휋푑2 휋푑2 푛 표 < 푛 표 4 4

2 2 푑푛 > 푛표푑표

2 2 푑푛 > √푛표푑표 = √6 × 0.003 = 0.00735푚 = 7.35푚푚

푑푛 > 7.35푚푚

푑푛 = 10푚푚 푤𝑖푙푙 푏푒 푢푠푒푑 (reasonable estimate)

Fig 4.5: Distributor plate and nozzles

ix) Determination of jet height 푙푗 From equation 24 [24], 0.3 2 0.2 퐿푢푝 휌푔푑표 푢 = 5.2 [ ] [1.3 ( 표 ) − 1] 푑표 휌푝푑푝 푔푑표

0.3 2 0.2 퐿푢푝 1.15 × 0.003 26.39 = 5.2 [ ] [1.3( ) − 1] = 16.24 0.003 532.72 × 0.0002 9.81 × 0.003

퐿푢푝 = 0.0487푚 = 48.7푚푚

From equation 23 [31, 32],

퐿푢푝~2퐿ℎ표푟

퐿푢푝 퐿 ≈ = 24.35푚푚 ℎ표푟 2 Horizontal nozzle jet configuration prevents particle flowback during shut down of reactor.

The outermost nozzles must be located at least 24.35mm from the walls of the reactor for the design to be considered safe.

x) Minimum Nozzle height From equation 25 [24], 푑 − 푑 푙 = 푛 표 푚 0.193 0.01 − 0.003 푙 = = 0.03627푚 ≈ 36푚푚 푚 0.193

A nozzle height of 67.5mm will be used.

xi) Thickness of the distributor plate, 푡푝 From the formula for calculating the maximum stress for a circular plate, with edges simply supported, under uniform loading, as given in equation 26), the thickness required for the plate can be calculated from

2 3(3+푣)푃푝푙푎푡푒 푟 푡푝 = √ 34) 8휎푚

6 2 From equation 4), Maximum allowable stress 푆 = 휎푚 = 52.5 × 10 푁/푚

푀푎푥푖푚푢푚 퐹표푟푐푒 푎푐푡푖푛푔 표푛 푡ℎ푒 푑푖푠푡푟푖푏푢푡표푟 푝푙푎푡푒 푃 = 35) 푝푙푎푡푒 퐴푟푒푎 표푓 푝푙푎푡푒

푚푔 푡표푡푎푙 푚푎푠푠 표푓 푏푖표푚푎푠푠 푎푛푑 푠푎푛푑×푎푐푐푒푙푒푟푎푡푖표푛 푑푢푒 푡표 푔푟푎푣푖푡푦 푃 = = 36) 푝푙푎푡푒 퐴 퐴푟푒푎 표푓 푝푙푎푡푒

푀푎푥𝑖푚푢푚 푓표푟푐푒 푎푐푡𝑖푛푔 표푛 푡ℎ푒 푑𝑖푠푡푟푢푏푢푡표푟 푝푙푎푡푒 = 푚푔 37)

From equation 32), 풎 = ퟒퟏ. ퟖퟒ풌품

풎품 = 41.84 × 9.81 = ퟒퟏퟎ.ퟒퟓ푵

From equation 35),

41.84 × 9.81 푃 = = 2094.13푁/푚2 푝푙푎푡푒 0.196 From equation 34),

3(3 + 0.24) × 2094.13 × 0.252 푡 = √ = 0.00174푚 = 1.74푚푚 푝 8 × 52.5 × 106

3 ′′ Assumed corrosion allowance = = 4.7625푚푚 16

Minimum Plate thickness with corrosion allowance, 푡푝 = 1.74푚푚 + 4.7625푚푚 = 6.5025푚푚

∴ 푡푝 ≈ 7.0푚푚

xii) Distributor plate nozzle thickness, 푡푛 To obtain the distributor plate nozzle thickness, the thicknesses derived from the hoop and longitudinal stress equations will be calculated. The greater value will be used as the nozzle thickness.

For Hoop Stress (from equation 2),

푃푑 106 × 0.01 푡 = 푛 = = 0.001618푚 푛 2푆퐸 − 1.2푃 2 × 52.5 × 106 × 0.6 − 1.2 × 106 = 0.1618푚푚

For Longitudinal stress (from equation 3),

푃푑 106 × 0.01 푡 = 푛 = = 0.000079푚 = 0.079푚푚 푛 4푆퐸 + 0.8푃 4 × 52.5 × 106 × 0.6 + 0.8 × 106

The minimum nozzle thickness of 0.1618푚푚 will be used. Since Stainless Steel 1 ′′ AISI 304L is resistant to corrosion, a low corrosion resistant value of = 16 1.5875푚푚

will also be used.

∴ Nozzle thickness is 0.1618 + 1.5875 = 1.7493푚푚 ≈ 1.8푚푚

4.3.3 Design of the Plenum or Windbox

The plenum also known as the wind box is located below the distributor plate. The major function of this section is the uniform pre-distribution of the fluidizing gas to all the nozzles of the distributor plate. The plenum is usually filled with packing materials to achieve uniform distribution of the fluidizing gas [33]. Several designs are usually employed for the plenum chamber. The preference for any design depends on the gas entry location into the windbox [34]. Litz [35] proposed designed criteria (table below) for the windbox depending on the orientation of the gas entry (i.e. Vertical or Horizontal)

Table 4.4: Plenum Design Criteria and Formula as proposed by Litz [35]

Orientation Criteria Formula Horizontal 퐷 ℎ = 0.2퐷 + 0.5푑 푑 > 푝푙푒푛푢푚 푒푛푡푟푦 푒푛푡푟푦 100 퐷 ℎ = 18푑 푑 < 푝푙푒푛푢푚 푒푛푡푟푦 푒푛푡푟푦 100 Vertical 퐷 ℎ = 3(퐷 − 푑 ) 푑 > 푝푙푒푛푢푚 푒푛푡푟푦 푒푛푡푟푦 36 퐷 ℎ = 100푑 푑 < 푝푙푒푛푢푚 푒푛푡푟푦 푒푛푡푟푦 36

The plenum chamber can be sectioned from the rest of the fluidized bed as evident in most commercial fluidized bed reactors. This eliminates the stagnant zone by increasing the pressure at the stagnant zone section at the expense of pressure in other sections of the reactor. Sectioning of the plenum, in general, complicates the design [33].

Using the correlation for horizontal orientation with 퐷 푑 > and appropriate value(s) in Table 4.2 (D=500mm and 푑 = 3.07 𝑖푛푐ℎ = 푒푛푡푟푦 100 푒푛푡푟푦 78푚푚 (퐷 푎푛푑 푑푒푛푡푟푦 푎푟푒 푟푒푎푠표푛푎푏푙푦 퐴푠푠푢푚푚푒푑). Horizontal orientation was chosen over vertical configuration because pressure loss is minimized.

퐻푝푙푒푛푢푚 = 0.2퐷 + 0.5푑푒푛푡푟푦 34)

= 0.2(500) + 0.5(78) ≈ 139푚푚

Chapter 5: FEA Analysis

Finite Element Analysis is a reliable method for testing the feasibility of a design before manufacturing. It offers a cost saving method of designing and analysis the various stresses, displacements and strain acting on each design member. The accuracy of the FEA results is dependent on the accuracy of the input data (geometry and boundary conditions).

5.1 Final Reactor Geometry Solidworks is used to sketch the geometry in 3D.

Fig 5.1: Fluidized Bed Reactor (Detailed Dimensioning of parts shown in Appendix)

Due to the complexity of the model, and the limitations in the computing power of ANSYS Student version, the FEA simulation will be done on each the main parts of the reactor instead of the entire assembly. The Solidworks model is converted to Parasolid and imported to ANSYS.

5.2 Meshing

The meshing of the entire reactor will be performed. Automatic meshing will be performed with the relevance center set to “Fine”. The Element Size will be set to a very small value suitable for running on ANSYS Student version without compromising the accuracy of the overall results

5.3 Boundary Conditions

Boundary conditions were carefully applied to match the different forces, pressure and supports that actually occurs. A thermal boundary condition is also applied to mirror a worst case scenario where the entire reactor is subjected to heating. The thermal boundary condition greatly affected the behavior of the materials used for this analysis.

5.4 Results

The results for the FEA simulation for the entire geometry will be analyzed. a) Geometry

Fig 5.2 Geometry of reactor

b) Meshing

A triangular mesh was used. Details of the mesh is shown in the table below. The mesh was applied over the entire geometry.

Fig 5.3 Mesh of Reactor Geometry

Object Name Mesh State Solved Display Display Style Body Color Defaults Physics Preference Mechanical Relevance 0 Element Order Program Controlled Sizing Size Function Adaptive Relevance Center Fine Element Size 7.0 mm Initial Size Seed Assembly Transition Fast Span Angle Center Coarse Automatic Mesh Based Defeaturing On Defeature Size Default Minimum Edge Length 1.7483e-002 mm Quality Check Mesh Quality Yes, Errors Error Limits Standard Mechanical Target Quality Default (0.050000) Smoothing Medium Mesh Metric None Inflation Use Automatic Inflation None Inflation Option Smooth Transition Transition Ratio 0.272 Maximum Layers 5 Growth Rate 1.2 Inflation Algorithm Pre View Advanced Options No Advanced Number of CPUs for Parallel Part Program Controlled Meshing Straight Sided Elements No Number of Retries Default (4) Rigid Body Behavior Dimensionally Reduced Mesh Morphing Disabled Triangle Surface Mesher Program Controlled Topology Checking No Pinch Tolerance Please Define Generate Pinch on Refresh No Statistics Nodes 3242946 Elements 1783260 Table 5.1: Reactor Geometry Meshing Details

54 c) Boundary Conditions

The boundary conditions applied to the reactor are as follows:

 Fixed Support at the base of the three legs  Pressure of 0.92MPa- About 9 times the atmospheric pressure  Thermal boundary condition of 1500C applied on the entire geometry  Force of 450N acting on the distributor plate due the biomass

Fig 5.4) Reactor Geometry showing boundary conditions (Pressure, Force and Fixed Support)

Fig 5.5) Section of Reactor geometry showing boundary conditions (Pressure, Force and Fixed Support)

Fig 5.6) Reactor Geometry showing thermal boundary condition

d) Results

Results for the Total Deformation, Equivalent Stress and Equivalent Elastic Strain will be analyzed. The deformation has to be very small so as not to affect the operation and stability of the reactor. As a rule, the Maximum Equivalent Stress must be less than the yield strength of the material in order to prevent material from going beyond its elastic limit.

i) Total Deformation

Fig 5.7) Total Deformation of the Reactor under operation

Fig 5.8) Section through the reactor showing Total deformation

ii) Equivalent (von-Mises) Stress

Fig 5.9) Equivalent (von-Mises) Stress simulation result for the Reactor

Fig 5.10) Section through the reactor showing Equivalent (von-Mises) Stress simulation result

iii) Equivalent Elastic Strain

Fig 5.11) Section through the reactor showing Equivalent Elastic Strain simulation result

From the above results for Total deformation, Equivalent Stress and Strain for the reactor, the following was observed

 The maximum total deformation of 0.40545mm (Fig 5.7 and 5.8) occurred at the center of the distributor plate and the gasket. This value is largely due to the thermal expansion of the A304L Stainless Steel at 1500C. The weight of the biomass had negligible effect on the deformation. This value is small and will likely not affect the overall operation of the reactor when in use. The simulation was done for a worst case scenario where the hot gas that enters the reactor at 400-500C and triggers an exothermic reaction of the biomass thus exceeding the inlet temperature. However, the temperature of the reaction in its worst case will not be up to 1500C so the maximum deflection will always be smaller than this value. The actual reaction temperature will be shown in the CFD simulation in the next chapter.  The maximum equivalent stress of 182MPa (Fig 5.9 and 5.10) occurs at the joint of the reactor and the stand. This value is less than the yield strength of A304L Stainless Steel

(which is about 210MPa). This implies that the stress value of 182MPa is within the elastic limit of the stainless steel material. Therefore, during operation the stainless steel can withstand the maximum stress without yielding.  The maximum equivalent stain occurs on the gasket. This is the lightest materials and it is expected to experience the most elongation (strain).

5.5 Results Verification/Validation

The validation of the results is essential in determining the accuracy of the result. The following were evaluated to validate the results

i) Boundary condition: The boundary conditions were accurate in describing the model as shown in Fig 5.5 and 5.6. ii) Results: The weight of the biomass acted downward on the distributor plate. This was observed in the Total Deformation results in Fig 5.7 and 5.8. The maximum total deformation of the distributor plate occurred at the center. This makes sense since the edges of the distributor plate are fixed. The maximum stress occurred at the leg stand. This is also accurate since the entire weight of the reactor were supported by the leg stand which has a very small cross sectional area relative to the reactor. The total deformation plot Fig 5.7 and 5.8 for the reactor was minimum at the legs because there were assumed to be fixed. These show the boundary conditions and results were accurate.

Chapter 6: Reaction Kinetics of Biomass Pyrolysis, CFD Modelling and Product Prediction

6.1 Reaction Kinetics of Biomass Pyrolysis

Knowledge of the reaction kinetics of biomass pyrolysis is important for prediction and optimization of pyrolysis products. Obtaining reliable kinetic rate data that can be used for a wide range of applications is very difficult. Kinetic models of pyrolysis of biomass can be broadly classified into three (Blasi, 1993): [2]

a) One stage global single reaction The overall reaction is considered as biomass → volatiles and char The rate of pyrolysis is dependent on the unpyrolyzed mass of biomass. The

decomposition rate of mass, 푚푏 in the primary pyrolysis process is written as 푑푚 푏 = −푘(푚 − 푚 ) 푑푡 푏 푐 −1 k is the first-order reaction rate constant (푠 ), 푚푐 is the mass of remaining char after complete conversion (kg), and t is the time in seconds [2].

b) One-stage, multiple reactions This involves several parallel reactions for describing the conversion of biomass into char and several gases. A one-stage simplified kinetic model is used to describe the different parallel reactions. This model is very useful in determining the product distribution [2].

Fig 6.1: One stage multiple reactions

c) Two-stage semi-global reaction This model involves both primary and secondary reactions which are occurring in series [2].

Fig 6.2 Two-stage semi-global reaction (by Shafizadeh et al)

6.2 CFD Modelling and Product Prediction of Biomass Pyrolysis

CFD remains an indispensable tool in the modelling of biomass fast pyrolysis in a fluidized bed reac tor. In this project, the Eulerian-Eulerian multiphase formulation is used to model the complex flow in the fluidized bed reactor. Two-fluid phases are applied in this computation- one for the reactant (biomass) and the other for the product (biochar, bio-oil and biogas). One stage, multiple reactions is used for this computation because the scope of this research involves only primary reactions. In this model, the biomass undergoes multiple reaction to produce three main products: biochar, biogas (non-condensable gases) and bio-oil (high molecular weight organic liquid) [36]. The results are validated by comparing with known experimental values. The commercial software ANSYS R19.0 was used to model the pyrolysis of the biomass in the reactor.

6.2.1 Methodology i) Geometry and Boundary Conditions

In every computer modelling, simplifying the geometry is a necessary step to minimizing the complexity of the computation. Because of this, a 2D representation of the model will be carried out. A 2D section at the middle of the reactor is used for this model. The section will contain the inlet (nozzles) and the outlet. The section will only be representative of the reaction zone (shell) of the fluidized bed reactor. The geometry dimensions were consistent with the dimensions of the 3D CAD model developed with Solidworks. The geometry was designed with ANSYS R19.0 Design Modeler.

Fig 6.3: Geometry of the 2D section of the reactor shell designed with ANSYS R19.0 DesignModeler

ii)Geometry Boundary Conditions

Three boundary conditions were used: inlet (on the nozzles), wall (for the sides of the reactor shell) and outlet. The boundary conditions are depicted below:

Fig 6.4: Inlet Boundary condition (on the sides of the nozzles)

Fig 6.5: Wall boundary condition

Fig 6.6: Outlet boundary condition

iii) Meshing

A triangular mesh was used in the geometry. An edge sizing of 1mm was used around the nozzle inlet for better accuracy in representing the boundary condition.

Fig 6.7: Geometry with meshing applied

The above geometry and boundary conditions are transferred to FLUENT for CFD analysis.

iv) CFD Inputs for ANSYS FLUENT R19.0 Analysis

 General- Pressure based solver, transient simulation with a gravitational acceleration of -9.81m/s2 applied on the y-axis.

 Models- Eulerian Multiphase (2 Eulerian phases) Phases assumed to be “mix’’ as the primary phase, mxture-template as the phase material; and biomass as the secondary phase . Formulation for volume fraction parameters is implicit. Interaction of both phases (mix and biomass): Drag Coefficient is Morsi- Alexander. Lift Coefficient is Moraga. No Turbulent Dispersion and Interaction applied. Restitution coefficient is 0.9. Heat transfer coefficient is Ranz-marshall. Heterogeneous Stiff Chemistry Solver is Applied.

- Energy is switched on. - Viscous Model is k-epsilon with Standard wall functions. - Species Transport with volumetric reaction, inlet Diffusion and Diffusion Energy Source applied. Phase is “mix’’ and Phase material is mixture- template.  Materials - The materials used in the simulation are the reactant (biomass) and products (biogas, biooil and char). Each of the material is treated as a fluid, hence the name fluidized bed. The properties of the materials (density, specific heat and thermal conductivity) must be known and inputted into the software. For the mixture- template, under ‘reaction’, the three competing reactions and their kinetic properties (Pre-exponential factor and Activation Energy) are also inputted. For this simulation, the kinetic properties used are consistent with the values given by Thurner et al [37] and is shown in Table 6.1

Pre-exponential Activation Energy(J/kg) factor(1/s) 푏𝑖표푚푎푠푠 → 푏𝑖표푔푎푠 144,000 88,600 푏𝑖표푚푎푠푠 → 푏𝑖표표𝑖푙 4,130,000 112,700 푏𝑖표푚푎푠푠 → 푏𝑖표푐ℎ푎푟 738,000 106,500 Table 6.1: Biomass Pyrolysis kinetic properties by Thurner et al [37]

 Boundary Conditions – For the inlet boundary condition, the Supersonic/Initial Gauge Pressure for the mixture phase is 101325 Pa, the ‘mix’ phase is set to the fluidizing velocity of 26.39m/s, temperature is set to 500C, and species is set to 1. The outlet gauge pressure for the mixture phase is 101325 Pa. All other boundary conditions properties’ are left at the default settings.  Initialization – The simulation is computed/initialized from the ‘inlet’. Markings (under Region Adaptation) are applied to indicate the volume fraction of biomass and temperature of the ‘mix’ phase at time of 0s. Two separate markings are made with coordinates of X Min(m)=0; X Max(m) = 0.5; Y Min (min)=0; Y Max(min)=0.8 and X Min(m)=0; X Max(m) = 0.5; Y Min (min)=0; Y Max(min)=1.2. Patching is applied to the above coordinates. The first set of

coordinates is patched for the ‘biomass’ phase for volume fraction of 0.8 while the second set of coordinates is patched for the ‘mix’ phase for temperature of 30oC.

 Run Calculation – The time step size is set to 0.005s and the number of time steps set to 2400. The maximum iterations/time step reporting interval is set to 20. The solution is then calculated.

6.3 Results

The simulation was run for 15 seconds and the results evaluated. The pressure, velocity and temperature contours for the simulation is shown below:

Fig 6.8a: Contour of Total pressure acting in the reactor due to fluidization

Fig 6.8b: Magnified view of Total Pressure contour

From Fig 6.8a &b, the pressure is maximum around the nozzle due to the exit of the fluidizing gas from the nozzle into the reactor. This pressure comes in direct contact with the particles in the reactor thus reducing its intensity. The pressure is however able to fluidize the particles. The somewhat round profile that forms around the top of the nozzle is consistent with the desired expectation. The individual nozzle gas exit port creates a path as it rises through the particles in the reactor. These individual gas paths grow as they join with neighboring paths and makes way to the exit of the reactor.

Fig 6.9: Contour of velocity magnitude of particles in the reactor

The above velocity plot is consistent with the pressure plot of Fig 6.8a &b. The velocity is highest at the nozzles exit ports, where the gas enters the reactor. The gas creates a pattern as it moves towards the exit of the reactor. The y-velocity plot below is necessary for understanding the velocity pattern of the flow in the reactor.

Fig 6.10: Contour of y-velocity of the particles in the reactor

The y-velocity plot above is consistent with the pressure and velocity plot of Fig 6.8 and 6.9. The above y-velocity plot is consistent with the expected velocity profile. There is a positive and negative velocity which indicates upward and downward motion respectively. The upward motion can be observed along the path of the gas flow while the downward motion is necessary to balance the effect of the upward motion of particles to ensure that there is uniform circulation and temperature distribution within the particles. While the upward motion is seen beside the nozzle, the downward motion can be seen around the top of the nozzles.

Fig 6.11: Contour of Static Temperature in the reactor

The temperature plot can be shown above. The maximum temperature in the reactor exceeds the fluidizing gas inlet temperature. This is possible because the pyrolysis reaction of biomass is exothermic. Due to the gas flow developing in the right side of the reactor as indicated by the velocity and pressure contour plots described below, most of the exothermic reaction occurs towards the right side of the reactor. Hence, the reason for increased temperature observed at the right side of the reactor.

Fig 6.12: Plot showing the pyrolysis products formation as a function of time

The product distribution of pyrolysis is a function of several factors including pyrolysis temperature, residence time of biomass particles, kinetic constants (dependent on type of biomass) etc. The different factors affecting yield has to be adjusted to achieve the desired product yield. For this simulation, for a biomass residence time of 15 seconds, gas inlet temperature of 500oC, and kinetic constants consistent with Thurner et al [37], the following result was obtained as shown in Fig 6.12.

Table 6.2: Biomass Pyrolysis Result Product Mass fraction Percentage Biogas 0.092233 ≈ 9.2% Bio-oil 0.643117 ≈ 64.3% Biochar 0.264586 ≈ 26.5%

6.4 Verification/Validation of Results

The biomass pyrolysis results obtained from this research was validated with experimental result obtained by Kalgo et al [38] as described by Panneerselvam et al [36]. The simulation parameters and biomass properties were fairly similar.

Table 6.3: Comparism of Research result with known experimental result Percentage Product Research Project Kalgo et al. Biogas 9.2% 13.56% Bio-oil 64.3% 63.15% Biochar 26.5% 23.29%

The values shown in Table 6.3 shows that the pyrolysis products obtained in this research project is close to the experimental values of Kalgo et al [38]. The slight difference might be due to different kinetic properties used due to different type of biomass accounted for. The results of this research is therefore valid.

7.0 Cost Analysis

In this chapter, an engineering bill of materials featuring the components of this design project is presented. SOLIDWORKS COSTING is used to give an estimate of each part used in the assembly. The cost shown gives an estimated value of the cost of materials and the manufacturing processes. The cost presented below will most likely be greater than the actual cost because the software used, evaluated each part to be produced by manufacturing processes applied to a block of material. This process is very wasteful, as it does not consider simple and less expensive manufacturing processes like welding of parts together.

Table 7.1: Engineering Bill of Materials Est. Estimated Item Description Part Number Quantity Weight Cost ($) (kg) 1 Reactor Shell 10001 1 500 20,000.00 Reactor Shell Outlet 2 10002 1 20 1,200.00 Cover 3 Top Cover 10003 1 400 15,388.26

4 Plenum and Base 10004 1 89 5,600.00 Studs for Top Cover and 5 10005 8 0.664 62.84 Shell 6 Flange 10006 2 394 2,000.00

7 Stub End 10007 2 90 1,000.00

8 Bolts 10008 24 58 240.00

9 Nuts 10009 24 14.26 216.00

10 Washers 10010 24 4.14 15.00

11 Wing nuts 10011 8 1.088 118.00 Hook for Shell Outlet 12 10012 1 0.336 22.84 cover 13 Distributor plate 10013 1 40.810 2,250.00 1,612.30 Total $48,112.94 kg

8.0 Conclusion

A fluidized bed reactor for biomass pyrolysis has been designed and validated. This reactor is of a smaller and more compact size compared to existing designs. This reactor is able to process biomass weighing up to 450N in 15 seconds producing biogas, biooil and biochar in the percentages 9.2%, 64.3% and 26.5% respectively. Process and design parameters can be varied to give product ratio in desired proportion. Commercialization of this design in a large scale will help in reducing the amount of carbon that would have been released into the atmosphere if the biomass were allowed to decay thus reducing global warming.

9.0 Future Directions a) Optimization of design parameters to maximize product yield b) Design/Development of a drive train, engine intallation and accessories for reactor mobility c) Design of a heating mechanism for supplying heat to the reactor through the combustion of the recycled biogas d) Continuous redesign for weight and cost reduction

10.0 References

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Appendix

2 1 640 500

55 32 8*10 37.35 B 4*TRUE R10 9.41 B

272 1117 22.89 313.04 313.04 19.80 TRUE R14.04 112 169.04 TRUE R20

159.52 84 152.82

251 202 12 83.34 UNLESS OTHERWISE SPECIFIED: NAME DATE

DIMENSIONS ARE IN MM DRAWN AATOLERANCES: FRACTIONAL CHECKED TITLE: ANGULAR: MACH BEND ENG APPR. TWO PLACE DECIMAL THREE PLACE DECIMAL Reactor Shell MFG APPR.

INTERPRET GEOMETRIC Q.A. PROPRIETARY AND CONFIDENTIAL TOLERANCING PER: COMMENTS: THE INFORMATION CONTAINED IN THIS MATERIAL DRAWING IS THE SOLE PROPERTY OF SIZE DWG. NO. REV . ANY REPRODUCTION IN PART OR AS A WHOLE FINISH WITHOUT THE WRITTEN PERMISSION OF NEXT ASSY USED ON A 10001 IS PROHIBITED. APPLICATION DO NOT SCALE DRAWING SCALE: 1:20 WEIGHT: SHEET 1 OF 1 2 1 2 1

458.54 R10.30 418.54 R8.30 36 60 R3 B 10.75 B 4 37.78

556.54

418.54 10 10

6 313.04 96.02 96.02 313.04 307.04 169.04 4 178.79

4*R20 4*R10 556.54 209.04 562.54

UNLESS OTHERWISE SPECIFIED: NAME DATE

DIMENSIONS ARE IN MM DRAWN AATOLERANCES: FRACTIONAL CHECKED TITLE: ANGULAR: MACH BEND ENG APPR. TWO PLACE DECIMAL Reactor Shell Outlet THREE PLACE DECIMAL MFG APPR. INTERPRET GEOMETRIC Q.A. Cover PROPRIETARY AND CONFIDENTIAL TOLERANCING PER: COMMENTS: THE INFORMATION CONTAINED IN THIS MATERIAL DRAWING IS THE SOLE PROPERTY OF SIZE DWG. NO. REV . ANY REPRODUCTION IN PART OR AS A WHOLE FINISH WITHOUT THE WRITTEN PERMISSION OF NEXT ASSY USED ON A 10002 IS PROHIBITED. APPLICATION DO NOT SCALE DRAWING SCALE: 1:5 WEIGHT: SHEET 1 OF 1 2 1 2 1

12.62 149.21 8.62

B 60.24 134.68 B 499.90

R320

TRUE R249.95 R335.09 8* 15 8* 33.18 42 8*TRUE R16.59 640 42

14.88

2*TRUE R7.50 499.90

UNLESS OTHERWISE SPECIFIED: NAME DATE

DIMENSIONS ARE IN MM DRAWN AATOLERANCES: FRACTIONAL CHECKED TITLE: ANGULAR: MACH BEND ENG APPR. TWO PLACE DECIMAL THREE PLACE DECIMAL Top Cover MFG APPR.

INTERPRET GEOMETRIC Q.A. PROPRIETARY AND CONFIDENTIAL TOLERANCING PER: COMMENTS: THE INFORMATION CONTAINED IN THIS MATERIAL DRAWING IS THE SOLE PROPERTY OF SIZE DWG. NO. REV . ANY REPRODUCTION IN PART OR AS A WHOLE FINISH WITHOUT THE WRITTEN PERMISSION OF NEXT ASSY USED ON A 10003 IS PROHIBITED. APPLICATION DO NOT SCALE DRAWING SCALE: 1:10 WEIGHT: SHEET 1 OF 1 2 1 2 1

TRUE R250

R250 R305.23

545.76 TRUE R250 TRUE R43 B 18* 20 R281.71 B

TRUE R75 TRUE R25

235.53 545.76

113.66 86 113.66

126.71 50 150 86 UNLESS OTHERWISE SPECIFIED: NAME DATE

DIMENSIONS ARE IN MM DRAWN AATOLERANCES: FRACTIONAL CHECKED TITLE: ANGULAR: MACH BEND ENG APPR. TWO PLACE DECIMAL Plenum(Windbox) THREE PLACE DECIMAL MFG APPR. INTERPRET GEOMETRIC Q.A. and base PROPRIETARY AND CONFIDENTIAL TOLERANCING PER: COMMENTS: THE INFORMATION CONTAINED IN THIS MATERIAL DRAWING IS THE SOLE PROPERTY OF SIZE DWG. NO. REV . ANY REPRODUCTION IN PART OR AS A WHOLE FINISH WITHOUT THE WRITTEN PERMISSION OF NEXT ASSY USED ON A 10004 IS PROHIBITED. APPLICATION DO NOT SCALE DRAWING SCALE: 1:10 WEIGHT: SHEET 1 OF 1 2 1 2 1

120

9 6.62 B B

9 6.62

9 9

115.50

UNLESS OTHERWISE SPECIFIED: NAME DATE

DIMENSIONS ARE IN MM DRAWN AATOLERANCES: FRACTIONAL CHECKED TITLE: ANGULAR: MACH BEND ENG APPR. TWO PLACE DECIMAL THREE PLACE DECIMAL Stud for Cover MFG APPR.

INTERPRET GEOMETRIC Q.A. PROPRIETARY AND CONFIDENTIAL TOLERANCING PER: and Shell COMMENTS: THE INFORMATION CONTAINED IN THIS MATERIAL DRAWING IS THE SOLE PROPERTY OF SIZE DWG. NO. REV . ANY REPRODUCTION IN PART OR AS A WHOLE FINISH WITHOUT THE WRITTEN PERMISSION OF NEXT ASSY USED ON A 10005 IS PROHIBITED. APPLICATION DO NOT SCALE DRAWING SCALE: 1:2 WEIGHT: SHEET 1 OF 1 2 1 2 1 R15

B 30 B 3

9 13.50

UNLESS OTHERWISE SPECIFIED: NAME DATE

DIMENSIONS ARE IN MM DRAWN AATOLERANCES: FRACTIONAL CHECKED TITLE: ANGULAR: MACH BEND ENG APPR. TWO PLACE DECIMAL THREE PLACE DECIMAL Wing nut MFG APPR.

INTERPRET GEOMETRIC Q.A. PROPRIETARY AND CONFIDENTIAL TOLERANCING PER: COMMENTS: THE INFORMATION CONTAINED IN THIS MATERIAL DRAWING IS THE SOLE PROPERTY OF SIZE DWG. NO. REV . ANY REPRODUCTION IN PART OR AS A WHOLE FINISH WITHOUT THE WRITTEN PERMISSION OF NEXT ASSY USED ON A 10011 IS PROHIBITED. APPLICATION DO NOT SCALE DRAWING SCALE: 1:1 WEIGHT: SHEET 1 OF 1 2 1 2 1

11.50 B 11.50 B

340 340

10 26.89

10 4 12.87 18

UNLESS OTHERWISE SPECIFIED: NAME DATE

DIMENSIONS ARE IN MM DRAWN AATOLERANCES: FRACTIONAL CHECKED TITLE: ANGULAR: MACH BEND ENG APPR. TWO PLACE DECIMAL THREE PLACE DECIMAL Hook for Shell MFG APPR.

INTERPRET GEOMETRIC Q.A. PROPRIETARY AND CONFIDENTIAL TOLERANCING PER: Outlet cover COMMENTS: THE INFORMATION CONTAINED IN THIS MATERIAL DRAWING IS THE SOLE PROPERTY OF SIZE DWG. NO. REV . ANY REPRODUCTION IN PART OR AS A WHOLE FINISH WITHOUT THE WRITTEN PERMISSION OF NEXT ASSY USED ON A 10012 IS PROHIBITED. APPLICATION DO NOT SCALE DRAWING SCALE: 1:5 WEIGHT: SHEET 1 OF 1 2 1 2 1

12.62 149.21 8.62

B 60.24 134.68 B 499.90

R320

TRUE R249.95 R335.09 8* 15 8* 33.18 42 8*TRUE R16.59 640 42

14.88

2*TRUE R7.50 499.90

UNLESS OTHERWISE SPECIFIED: NAME DATE

DIMENSIONS ARE IN MM DRAWN AATOLERANCES: FRACTIONAL CHECKED TITLE: ANGULAR: MACH BEND ENG APPR. TWO PLACE DECIMAL THREE PLACE DECIMAL Top Cover MFG APPR.

INTERPRET GEOMETRIC Q.A. PROPRIETARY AND CONFIDENTIAL TOLERANCING PER: COMMENTS: THE INFORMATION CONTAINED IN THIS MATERIAL DRAWING IS THE SOLE PROPERTY OF SIZE DWG. NO. REV . ANY REPRODUCTION IN PART OR AS A WHOLE FINISH WITHOUT THE WRITTEN PERMISSION OF NEXT ASSY USED ON A 10003 IS PROHIBITED. APPLICATION DO NOT SCALE DRAWING SCALE: 1:10 WEIGHT: SHEET 1 OF 1 2 1

**Distributor plate nozzle dimension addendum. Circular pattern of 6 nozzles is cut-extruded from the nozzle