Calcination in an electrically heated bubbling fluidized bed applied in calcium looping

(1)

Faculty of Technology, Natural sciences and Maritime Sciences

Campus Porsgrunn FMH606 Master's Thesis 2020 Energy and Environmental Technology

Calcination in an electrically heated bubbling fluidized bed applied in calcium

looping

Nastaran Ahmadpour Samani

(2)

Course: FMH606 Master's Thesis, 2020

Title: Calcination in an electrically heated bubbling fluidized bed applied in calcium looping Number of pages: 149

Keywords: Electrically-heated calciner, Bubbling fluidized bed, design, Minimum fluidization velocity, Binary-particles system, Terminal settling velocity, Cement industry, Enhanced cement raw meal fluidization, Computational Particle and fluid Dynamics (CPFD), Barracuda®, simulation

Student: Nastaran Ahmadpour Samani

Supervisor:

Co-supervisor:

Lars-André Tokheim Chameera Jayarathna

External partner: SINTEF Tel-Tek (Chameera Jayarathna), Norcem (Tor Gautestad)

Availability: Open

(3)

The University of South-Eastern Norway takes no responsibility for the results and conclusions in this student report.

Summary:

Switching fossil fuels to green electricity as the energy source to decarbonate the raw meal in the calciner can eliminate the CO2 emissions produced through fuel combustion and also provide a basis for simple capture of the CO2 generated through calcination, as CO2

is the only gaseous product exiting from the electrified calciner. For this reason, an electrically-heated fluidized bed reactor was designed as a calciner and its applicability and cost estimation were carried out.

A mass and energy balance for steady-state conditions was conducted, so that relevant temperature, flow rates, and duties in the electrically-heated FB reactor and heat exchanger have been calculated by MATLAB code.

The key parameters of FB reactor such as minimum fluidization velocity, minimum bubbling velocity, terminal settling velocity, and the reaction time based on the particle size distribution were calculated. The fluidizability of the fine limestone particles was tested by a cold-bed BFB unit and it revealed that owing to the fine particle sizes of the raw meal, there are strong cohesive forces between the particles. Hence, a conventional bubbling fluidized bed is difficult to fluidize Geldart C particles. The identical system was simulated by Barracuda® and the results of the simulations had a good consistency with the experiments.

A binary-particle fluidization system, mixing fine powders with the coarse particles, was proposed to enhance the flowability of fine particles. The fluidized bed calciner process was designed as a semi-batch process operating in two modes; the calcination mode (with a low gas velocity) and the entrainment mode (with a higher velocity). After the raw meal particles have been calcined, they have to be separated from the coarse, inert particles.

This can be done by increasing the velocity of the CO2 used for fluidization to a value sufficiently high to entrain the raw meal particles, but still sufficiently low that the coarse, inert particles are not entrained. The inert particles may provide a homogeneous distribution of the fine particles and help to fluidize them. The aggregation and clustering of the fine particles will decrease due to collisions with inert coarse particles. The inert particles will also provide a thermal energy reservoir through their heat capacity and thereby contribute to a very stable bed temperature, which is advantageous in the control of the process.

The operational conditions at 1173 K, such as the particle size distribution of the inert particles and the fluidization gas velocity were calculated by the Barracuda simulations.

The inert particles with the diameter range of 550-800 µm and the velocities in the calcination and entrainment modes equal to 0.18 m/s and 3 m/s appeared as suitable for the calciner operation. The simulations showed that at the velocity of 0.18 m/s, 7.6% of fine particles may be entrained. However, by comparing the CO2 residence time with the reaction time of particles, it was concluded that all fine powders were calcined before leaving the bed.

(4)

A circular cross-sectional area was chosen to have the symmetric distribution of particles.

The design calculation results of horizontal and vertical arrangements of heating tubes showed that the pressure drop in the vertical layout of heating tubes is lower. The impact of the number of reactors on the important parameters of FB reactors illustrated that the capital cost of one reactor was about 50% less costly than two parallel reactors. Hence, one reactor with a diameter of 17.32 m and a height of 3.3 m in the vertical arrangement of heating tubes was suggested as an appropriate alternative.

The required CO2 gas for fluidization was compared with the produced CO2 through the calcination process and showed that enough CO2 is produced during the calcination process. Therefore, self-fluidization can occur in the phase of the calcination. There is no negative impact on the process and the quality of the product due to the replacement of the FB reactor instead of traditional calciners. Minor changes and modifications may happen. In order to produce 1 MT of clinker per year, 85 MW electrical energy is required.

The cost estimation of the FB reactor and the fan as a pressure compensation was conducted resulting in the annualized cost of 43 € per ton of CO2 captured. The equivalent annual costs for capital and operational expenditure were calculated 2.56 and 201.42 MNOK, respectively. The criterion for the material selection was high-temperature resistance. Stainless steel FB reactor with a layer of calcium aluminate refractory material to protect against hot gas and silicon carbide heating tubes were selected to manufacture of the reactor. Centrifugal radial fan was identified as equipment for pressure compensation.

(5)

Preface

This master’s thesis titled “Calcination in an electrically heated bubbling fluidized bed applied in calcium looping” was done at the University of South-Eastern Norway (USN), Porsgrunn.

It has been written to fulfill the graduation requirement of a Master of Science degree in Energy and Environmental Technology.

This master thesis has been defined in line with the collaboration of USN, SINTEF Tel-Tek, and Norcem in a research project “Combined calcination and CO2 capture in cement clinker production by use of the CO2-neutral electrical energy”.

The picture on the first page was taken from the website of Norcem, Brevik.

It was a highly valuable experience for me to be involved in this research project and to work under the close supervision of Prof. Lars André Tokheim.

First and foremost, I would like to express my sincere gratitude and appreciation to Prof. Lars André Tokheim for his great support throughout the project. He always was available when I needed help and I benefited a lot from his knowledge and guidance to carry out this study. I would also like to express gratitude to my co-supervisor Dr. Chameera Jayarathna for his support, valuable suggestions, and encouragement throughout the project. I would also like to express my appreciation to Tor Gautestad for consultation, encouragement, and interest. This thesis would not have been a proceeding without the help and collaboration of them.

I extend my gratitude to Prof. Britt Margrethe Emilie Moldestad and all Ph.D. fellows in her team to collaborate with me for accessing the Barracuda license. I would also like to thank the staff of the IT department, for providing me with various technical supports to access Solid Works and remote access to computers and software during the Covid-19 condition.

I would like to thank my husband Ali who always was there for me during this journey and encouraged me to do my best in completing this thesis.

Finally, I would like to thank my family for their full support during these two years to fulfil the master’s degree.

Porsgrunn, May 2020

Nastaran Ahmadpour Samani

(6)

Nomenclature

List of symbols

Symbol Description Unit

𝐴_𝐵𝐹𝐵 The cross-sectional area of the reactor [𝑚²]

𝐴_{𝑝,𝑝𝑟𝑜𝑗} Projected area of the particle [m²]

𝐶_𝑝_𝑒𝑓𝑓 Adjusted specific heat [ 𝐽

𝑘𝑔 𝐾]

𝐶_𝐷 Drag coefficient [-]

𝐸_{𝑔𝑒𝑛,𝑐𝑎𝑙} Generated energy [MW]

𝐸_{𝑖𝑛,𝑐𝑎𝑙} Inlet energy [MW]

𝐸_{𝑜𝑢𝑡,𝑐𝑎𝑙} Outlet energy [MW]

𝐹₄₅ The mass fraction of particles having the diameter of

particles smaller than 45 µm. [-]

𝐻_{ℎ𝑒𝑎𝑡} The height of the heating area [m]

𝐻_𝑏𝑒𝑑 The total height of the bed [m]

𝐻_𝑐𝑎𝑙 The enthalpy of calcination [MJ/kgCO2]

𝐻_{𝑜𝑡ℎ𝑒𝑟,𝑐𝑎𝑙} The enthalpy of other meal reaction [MJ/kgCO2]

𝐿_𝑒 The average length of the heating tubes [m]

𝑀_𝐶𝑂2 The molecular masses of CO2 [kg/mol]

𝑀_{𝐶𝑎𝐶𝑂3} The molecular masses of CaCO3 [kg/mol]

𝑅𝑒_𝑚𝑓 Reynolds number [-]

𝑇_𝐶𝑂₂_{,𝑜𝑢𝑡} The outlet temperature of CO2 [K]

𝑇_𝐶𝑂₂_{,𝑟𝑒𝑐𝑦} The temperature of recycled CO2 [K]

𝑇_𝑃𝐻𝑀 Preheated temperature [K]

𝑇_{𝑎𝑖𝑟,𝑖𝑛} Excess cooling air temperature [K]

𝑇_{𝑎𝑖𝑟,𝑜𝑢𝑡}

The outlet temperature of the airstream [K]

(10)

𝑇_{𝑖𝑛𝑒𝑟𝑡} The temperature of inert particles [K]

𝑇_𝑟𝑒𝑓 Reference temperature (𝑇𝑟𝑒𝑓=25 °C and pref= 1atm) [K]

𝑈_𝑚𝑏 Minimum bubbling velocity [m/s]

𝑐_{𝑝,𝐶𝑂}₂_{,𝑐𝑎𝑙}

Specific heat at the constant pressure of CO2 at 𝑇_𝑐𝑎𝑙 [J/kg K]

𝑐_{𝑝,𝐶𝑂}₂_{,𝑟𝑒𝑐𝑦} Specific heat at the constant pressure of recycled CO2

at 𝑇_𝐶𝑂₂_{,𝑟𝑒𝑐𝑦} [J/kg K]

𝑐_{𝑝,𝐶𝑂2,𝐻𝐸𝑋} Specific heat at the constant pressure of CO2 at the average temperature the hot side

[J/kg K]

𝑐_{𝑝,𝑃𝐻𝑀} Specific heat at the constant pressure of preheated meal at 𝑇_𝑃𝐻𝑀

[J/kg K]

𝑐_{𝑝,𝑎𝑖𝑟,𝐻𝐸𝑋} Specific heat at the constant pressure of air at the average temperature of the cold side.

[J/kg K]

𝑐_{𝑝,𝑚𝑒𝑎𝑙,𝑐𝑎𝑙} Specific heat at the constant pressure of calcined meal at 𝑇_𝑐𝑎𝑙

[J/kg K]

𝑐_𝑝 Specific heat [J/kg K]

𝑑_𝑒 The diameter of the heating tubes [m]

𝑑_𝑝 Particle diameter [m]

𝑓_{𝐸𝑄,𝐶𝑆} Equipment cost of carbon steel factor [-]

𝑓_{𝑃𝐼,𝐶𝑆} The factor of piping cost for carbon steel [-]

𝑓_{𝑇𝐼𝐶,𝐶𝑆}

The total installed factor for carbon steel [-]

𝑓_𝑇𝐼𝐶 The total installed cost factor [-]

𝑓_{𝑝𝑟𝑒𝑐𝑎𝑙} The calcination degree [-]

𝑘_𝑠 Allowed stretch tension [N/m²]

𝑚_𝑃𝐻𝑀 Mass of preheated meal [kg]

𝑚_𝑓 Material factor [-]

𝑛_𝑟𝑜𝑤 The number of horizontal rows [-]

𝑛_{𝑡𝑢𝑏𝑒} The total number of heating elements [-]

𝑛𝑡𝑢𝑏𝑒_𝑖𝑛_𝑟𝑜𝑤 The number of heating elements in a horizontal row [-]

(11)

𝑡_𝐶𝑂2 CO2 gas residence time [s]

𝑡_𝑝 Residence time of particles [s]

𝑣_𝑐𝑎𝑙 Velocity in calcination mode [m/s]

𝑣_𝑡 Terminal settling velocity [m/s]

𝜀_𝑚𝑓 Bed voidage at minimum fluidization [-]

𝜂_{𝑒𝑙,ℎ𝑒𝑎𝑡} Electricity to heat conversion efficiency [−]

𝜌_{𝑏𝑢𝑙𝑘} Bulk density [kg/m³]

𝜌_𝑝 Particle density [kg/m³]

𝜎_{𝑦𝑖𝑒𝑙𝑑} Yield strength [N/m²]

∆𝑃 The pressure drop [Pa]

FD The drag force [N]

ℎ Distance between adjacent rows [m]

Nup Nusselt number [-]

Pr Prandtl’s number [-]

t reaction Reaction time [s]

U Overall heat transfer coefficient [ ^𝑤

𝑚²𝐾]

Umf Minimum fluidization velocity [m/s]

Wt % Mass fraction of particles [-]

Δ𝑇_{𝐻𝐸𝑋,𝑚𝑖𝑛} The minimum temperature difference of heat exchanger

[K]

𝐴𝑟 Archimedes number [-]

𝐷 The diameter of the reactor [m]

𝐸𝑒𝑙, 𝑐𝑎𝑙, 𝑠𝑢𝑝𝑝𝑙𝑦 The required supply of electrical energy [MW]

𝐻 The height of the reactor [m]

𝑎 Distance between adjacent tubes [m]

𝑘 Thermal conductivity [W/m K]

𝑤𝐶𝑎𝐶𝑂3, PHM The CaCO3 content in the preheated meal [kg/kg]

(12)

𝜌 Density [kg/m³]

𝜑 Sphericity [-]

𝑄̇_ℎ The heat transfer for sensible heat [𝑊]

𝑉̇_𝐶𝑂₂ CO2 flow rate

[𝑚³ 𝑠 ] 𝑚̇_𝐶𝑂₂_{,𝑝𝑟𝑜} CO2 generated by the meal in the calciner [t/h]

𝑚̇_𝑃𝐻𝑀 The preheated meal rate [t/h]

𝑚̇_𝑐𝑜₂_{,𝑟𝑒𝑐𝑦} the CO2-recycled rate [t/h]

𝑚̇_{𝑚𝑒𝑎𝑙,𝑐𝑎𝑙} Calcined meal rate [t/h]

(13)

BFB Bubbling Fluidized Bed

CAPEX Capital Expenditure

CPFD Computational Particle Fluid Dynamics

CS Carbon steel

DFT Detailed Factor Table

HE Heat Exchanger

OPEX Operational Expenditure

PHM Preheated Meal

PPFB Particle Powder Fluidized Bed

PSD Particle Size Distribution

SS Stainless Steel

TIC Total Installed Cost

(14)

List of figures and tables

List of figures

Figure 2.1: Different fluidization regimes [10] ... 21

Figure 2.2: The Geldart classification of particles for air at ambient conditions [10] ... 22

Figure 2.3: Pressure drop vs superficial gas velocity [10] ... 25

Figure 2.4: Schematic diagram of the experimental apparatus [29] ... 28

Figure 2.5: A cement regular kiln process with two preheater strings [37] ... 31

Figure 2.6: A modified cement kiln process applying electrical energy for calcination [37] . 32 Figure 3.1: Process flow diagram for the specified system of project ... 33

Figure 3.2: The effect of the recycling rate for CO2 into the required supply of electrical energy and outlet CO2 temperature of the heat exchanger ... 39

Figure 3.3: Particle density vs bulk density [43] ... 43

Figure 3.4: Cumulative distribution of limestone particles ... 45

Figure 3.5: Comparison of particle size distribution by considering clustering factor and the exact values ... 46

Figure 3.6: Terminal settling velocity based on different particle sizes and gases... 47

Figure 3.7: Minimum fluidization velocity based on the particle diameters for CO2 and Air 49 Figure 3.8: Minimum bubbling velocity based on the particle size for CO2 and air ... 50

Figure 3.9: velocity window for a binary-particles system (CaCO3 and SiO2) with airflow... 51

Figure 3.10: Velocity window for a binary-particle system (CaCO3 and SiO2) with CO2 gas 52 Figure 3.11: velocity window by considering clustering factor ... 52

Figure 3.12: Reaction time for different particle exact sizes of limestone ... 53

Figure 4.1: Cold-flow lab-scale fluidized bed unit ... 54

Figure 4.2: cold-flow BFB experiments with different inlet gas flow rates ... 56

Figure 4.3: Experimental results of pressure drop vs superficial gas velocity ... 57

Figure 5.1: Phase A: before feeding limestone powders ... 59

Figure 5.2: Fluidization of inert particles ... 60

Figure 5.3: Phase B: feeding the preheated limestone powders ... 61

Figure 5.4: Phase C: heating the limestone particles from preheated temperature to calcination temperature ... 62

Figure 5.5: Different types of heat transfer to limestone particles ... 63

Figure 5.6: Phase D: Calcination of the batch of limestone particles ... 64

Figure 5.7: entrainment of calcined meal ... 65

(15)

Figure 6.2: (a) Original model geometry and (b) Grid and original CAD geometry together 68

Figure 6.3: The output of Grid check utility ... 68

Figure 6.4: initial conditions of a) single-particle system and b) binary-particle system, and c) Boundary conditions ... 70

Figure 6.5: a) transient data and b) flux plane ... 71

Figure 6.6: The pressure drop versus velocity for experimental and simulation results ... 72

Figure 6.7: Particle distribution at the velocity of 0.08 m/s... 73

Figure 6.8: The impact of clustering factor in the simulation results of limestone particles ... 74

Figure 6.9: Particle distribution in calcination mode at different periods at v=0.18 m/s ... 77

Figure 6.10: Volume fraction of fine powders and inert particles in different periods at v=0.18 m/s ... 78

Figure 6.11: The initial condition of simulation with the ratio 1/10 ... 79

Figure 6.12: Distribution of limestone and silica sand particles with the ratio 1/10 at different periods, v=0.18m/s ... 80

Figure 6.13: Volume fraction of fine and inert particles with ratio 1/10, v=0.18 m/s ... 81

Figure 6.14: The distribution of fine and coarse particles in the entrainment stage ... 83

Figure 7.1: A schematic of a BFB reactor with a horizontal layout of heating tubes ... 85

Figure 7.2: Design procedure of FB reactor with a horizontal arrangement of heating tubes . 87 Figure 7.3: Heat transfer to limestone particles for heating up to 900 C and calcination ... 88

Figure 7.4: Different dimensions of the reactor and heating tubes ... 90

Figure 7.5: The arrangement of heating tubes in a row ... 93

Figure 7.6: The effect of distance between adjacent heating tubes in a row, a, on a) the height of the reactor, b) number of tubes, c) CO2 residence time, d) particle residence time ... 98

Figure 7.7: Schematic of the vertical arrangement of heating tubes inside the reactor ... 102

Figure 7.8: Cross-section of the FB reactor with the vertical arrangement of the heating tubes ... 103

Figure 7.9: Design procedure of the FB reactor with the vertical arrangement of heating tubes ... 104

Figure 8.1: The section of the designed FB reactor ... 117

(16)

Table 2.1: The correlations to predict minimum fluidization velocity [14] ... 24

Table 2.2: Experimental conditions of tests done by Tashimo et al [29] ... 28

Table 3.1: Mass balance design basis values ... 34

Table 3.2: Energy balance design basis values ... 35

Table 3.3: Properties of applied gas in bubbling fluidized bed at different temperatures [47] 46 Table 6.1: The minimum fluidization velocity and terminal velocity of inert particles ... 75

Table 6.2: The values of entrained mass in calcination mode ... 76

Table 6.3: The value of mass entrained at 18 m/s with ratio 1:10 ... 79

Table 6.4: The values of the entrained mass of fine and inert particles in entrainment stage . 82 Table 7.1: Assumptions for the design of the FB reactor by the horizontal arrangement of heating tubes ... 90

Table 7.2: The impact of the number of reactors on the size of the reactor ... 101

Table 7.3: Design calculation results based on the types of the arrangement of heating tubes and the number of reactors ... 108

Table 8.1: Detailed factor table (DFT) [59] ... 110

Table 8.2: High-temperature alloys [61] ... 113

Table 8.3: Different Refractory materials [63] ... 114

Table 8.4: Allowed strength tension vs temperature [66] ... 115

Table 8.5: Equipment price from the cost estimation website [70] ... 117

Table 8.6: Currency converting ... 118

Table 8.7: Inflation from 2002 to 2020 [72] ... 118

Table 8.8: Updated cost data for facilities ... 118

Table 8.9: Total installed cost ... 119

Table 8.10: Total cost of the system ... 119

Table 8.11: Cost estimation for the case of having 2 parallel reactors ... 120

Table 8.12: The price of electricity in Norway [73] ... 121

(17)

1 Introduction

In this chapter, background, problem description, and objectives are discussed.

1.1 Background

Concrete is the most consumed man-made material in the world. Cement is a key constituent in concrete, contributing to 8% of global anthropogenic CO2 emission. Cement industries produce 4 billion tonnes of cement per year and cement is one of the main contributors to climate change[1].

Due to urbanization and emerging new economies especially in Asia, it is predicted that the demand for the cement will grow up by 12-23% by 2050 compared to the year 2014. Therefore, strict carbon capture mitigation strategies are needed in the cement industry to comply with the Paris agreement on climate change. For reducing the CO2 emission in the cement industry, some measures should be taken such as

1) applying CO2 capture methods,

2) improving existing plans to recuperate thermal and electric efficiency, 3) replacement of low-carbon fuels instead of fossil fuels, and

4) substitution of clinker as the largest carbon emitter [2-4].

In the cement process, two major stages are emitting the highest amount of carbon dioxide.

• The first stage is producing the major constituent of cement, clinker, during the calcination of limestone (CaCO3) into lime (CaO) and CO2, during the chemical reaction 𝐶𝑎𝐶𝑂₃ → 𝐶𝑎𝑂 + 𝐶𝑂₂. Thus, this process could emit 65% of CO2.

• The second part is fossil fuel combustion by producing 35% of CO2 emissions [5, 6].

1.2 Problem description

Switching fossil fuels to renewable sources in power plants and energy efficiency are important issues for having a green future. Therefore, utilizing green electricity to produce clinker and applying an appropriate carbon capture technology to the reduction of CO2 emissions of the calcination process can be a good solution to mitigate the emission of CO2 in the cement industry.

Utilizing green electricity instead of fossil fuels to decarbonate the raw meal in the cement kiln process can eliminate the CO2 emissions produced through fuel combustion and also provide a basis for simple capture of the CO2 generated through calcination, as CO2 is the only gaseous product exiting from the electrified calciner[6].

As a result of modifying the calciner, different reactors can be used as a calciner. In the current master thesis’ project, an electrically heated fluidized bed (FB) reactor is being developed to calcine the raw meal used in cement manufacturing. The FB will replace the traditional entrainment calciner used in many cement plants. The purpose is to enable efficient indirect heat transfer in the bubbling bed and hence obtain pure CO2 as the gaseous product from the calciner.

(18)

1.3 Objectives of the study

The main objective of the project is to design a fluidized bed (FB) reactor as a calciner in the kiln process and evaluate the applicability and cost of this concept.

In order to achieve the goal of the project, project objectives can be classified into:

1. Evaluating the FB reactor and its ability

2. Making a flow diagram for the appropriate case of simulation

3. Applying mass and energy balance equations to calculate temperatures, flow rates and duties

4. Making a process simulation of the FB and simulate several cases by changing key parameters of the design

5. Measuring the main characteristic of the design of BFB reactors, minimum fluidization velocity and pressure drop by experiment and through computational particle and fluid dynamics (CPFD) simulations applying the commercial software Barracuda®

6. Suggesting a FB reactor design and define the design basis values 7. Investigating the impacts of new calciner on the kiln system

8. Estimating the required size of different equipment such as FB reactor, and the CO2 fan 9. Estimating the investment cost (CAPEX) and operational cost (OPEX) of the

appropriate case per avoided CO2 unit (€/tCO2)

To fulfill the requirements and objectives of the project, answering some key questions are essential:

1. What is the concept of the fluidization phenomena?

2. What are the advantages of fluidized beds?

3. How BFB can transfer the heat generated by resistance heating to the meal?

4. What are the specifications and requirements of design a BFB?

5. How can the minimum fluidization velocity and terminal settling velocity be calculated based on the several particle sizes?

6. How can a potential limestone particle cohesion problem be dealt with?

7. How can an appropriate velocity range for the BFB reactor be determined?

8. How much CO2 should be recycled?

9. What is the value of the cross-sectional area of the reactor?

10. How many reactors do we need?

11. What are the values of temperatures, duties, and flow rates in the new system?

12. What are the major impacts on the kiln process system?

13. What are the required sizes of other equipment such as CO2 fan in the selected case?

14. What is the estimated cost of the new system? Is it economically beneficial compared to the commercially available carbon capture technologies?

The task description and Work Break down Structure are presented in appendix A and B respectively.

(19)

1.4 Organization of the report

This report contains 8 chapters. The first chapter describes the background of the study, problem description, and the objectives of the thesis. Chapter 2 gives an overview of the fluidization as well as the key parameters of the design of BFB reactors. The mass and energy balance calculations, minimum fluidization velocity, minimum bubbling velocity, and terminal settling velocity are discussed in Chapter 3. In the next chapter, the procedure and results of the experimental tests in a cold-flow lab-scale fluidized bed unit are presented. The enhanced raw meal fluidization system through mixing with coarse particles is explained in Chapter 5.

In Chapter 6, the development of the model in Barracuda® and the conduction of simulations for different cases will be explained step by step. Moreover, the simulation results for the experimental FB unit are presented and compared with the experimental results. Besides, Barracuda is applied to find the operational conditions at the temperature of 1173 K and the velocities of calcination and entrainment mode for the design of the binary-particle system.

Chapter 7 contains all procedures and design calculations of the FB reactor. The material selection and cost estimation are presented and discussed in Chapter 8.

(20)

2 Theory and literature review

This section starts with the general information of fluidization phenomenon and regimes, then some important parameters in fluidization are discussed.

2.1 Overview of fluidization

Fluidization is the process of quiescent solid particles caused to behave like a fluid by an upward stream of gas or liquid blown into the solid-filled reactor. Fluidization applications in a wide range of engineering operations can be classified into two groups;

• Physical operations consist of heating transfer, absorption, conveying system, fine powder mixing, etc. and

• Chemical operations consist of chemical reactions of gases on solid catalysts or chemical reactions of solids

One of the well-known methods in the process industry is the fluidized bed. Fluidized bed processes have commercially applied since the 1920s by emerging of the Winkler coal gasifier in Germany. Fluidized bed reactors have several advantages, such as high heat and mass transfer rate, good mixing properties, and close to isothermal conditions. Because of these advantages, the fluidization concept is widely used in different engineering applications such as the reforming of hydrocarbons, gasification, aluminium production, and calcination [7-10].

2.2 Fluidization regimes

Depending on the gas velocity as well as gas and solid characteristics, the fluidized bed behaves differently. As shown in Figure 2.1, there are several fluidization regimes.

Figure 2.1 (a) shows that the fluid penetrates the void spaces between the particles at a low flow rate; however, the height remains as the bed at rest. This step is a fixed bed. With increasing the superficial gas velocity, at a point, the drag force of upward flow gas becomes equal to the weight of the particles in the bed (Figure 2.1(b)). This point is the onset of fluidization and is one of the main characteristics in the fluidization, minimum fluidization velocity (Umf). In the gas-solid system, when superficial gas velocity increases above the minimum fluidization velocity, a smooth expansion of the bed can occur. The bed like in Figure 2.1 (c) can be observed only when fine particles have been exposed to high-pressure dense gas.

With increasing the superficial gas velocity beyond minimum fluidization velocity, the bubbles are formed and the coalescence of bubbles and formation of channelling appear. At this point, the bubbling fluidized bed occurs at a superficial gas velocity very slightly higher than the minimum fluidization velocity, this velocity so-called minimum bubbling velocity (Umb). The height of the bed does not increase much beyond the height in the minimum fluidization (Figure 2.1 (d)). This bed is bubbling fluidization and can take place under the special condition of the fluidization of very dense particles by fluids with low density [10, 11].

With increasing the velocity further, the coalescence of bubbles occurs and the bubbles in the fluidized bed integrate and rise. The size of the bubble may become as large as the diameter of the bed provided that the ratio of the height to the diameter of the bed is sufficient. This condition is the so-called slugging as it is illustrated in Figure 2.1 (e). If the velocity reaches

(21)

greater than the terminal settling velocity, the particles from the upper surface of the bed can eject from the bed and turbulent motion of solid clusters and voids of gas can be formed, like Figure 2.1(g), the bed is called turbulent beds. In the condition that the velocity is increasing further, the fluidized bed converts to an entrained bed with a dilute or lean phase fluidized bed leading to pneumatic transport of solids, as it is shown in Figure 2.1 (h). The majority of modern cement plant calciners today operate close to the regime (h) in figure [10, 11].

Figure 2.1: Different fluidization regimes [10]

2.3 Geldart classification of particles

Based on experimental data and empirical observations, Geldart [12] categorized particles into four types of particle behavior; A, B, C, and D, depending on particle size and density difference between solids and the fluidization gas, Figure 2.2.

• Group C: cohesive and very fine powders (𝑑̅_𝑝< 30𝜇𝑚). Due to the large inter- particular forces, they are extremely difficult to fluidize.

(22)

• Group A: aeratable particles (20 < 𝑑̅_𝑝 < 100𝜇𝑚). They have a small mean particle size and low particle density (<~1.4 g/cm³). They fluidize easily with smooth fluidization at low superficial gas velocities before bubbles appear.

• Group B: sand-like particles (40 < 𝑑̅_𝑝 < 500𝜇𝑚 and 1.4 < 𝜌_𝑠 < 4 𝑔/c𝑚³). they fluidize easily and once the velocity is more than minimum fluidization velocity, bubbles are formed.

• Group D: ‘spoutable’ particles (large or very dense particles). It is difficult to fluidize in deep beds. With increasing the gas velocity, a jet can be formed in the bed, and particles can act as exploding bubbles or spouting behaviour [10, 12, 13].

Geldart graph can be easily used for the air fluidization at ambient conditions and for 𝑢_° <

10 𝑢_𝑚𝑓.

Figure 2.2: The Geldart classification of particles for air at ambient conditions [10]

2.4 Minimum fluidization velocity

The minimum fluidization velocity (𝑈_mf) is defined as the superficial gas velocity when the weight of the solid particles is equal to the drag force of the upward gas flowing. At this point, all particles are suspended in the fluid on this is the onset of fluidization. The minimum fluidization velocity is one of the most important characteristics associated with the design of the fluidized bed reactor.

(23)

2.4.1 Theoretical method to determine minimum fluidization velocity

Minimum fluidization velocity can be calculated theoretically by applying some correlations based on the solid and fluid characteristics such as density, drag coefficient, and viscosity.

From 1950, there have been published several correlations to predict the 𝑈_mf [8, 10, 14]

In most cases, the minimum fluidization velocity is defined as a function of

• particle characteristics such as particle diameter (dp) and density (ρp) and particle sphericity (φ),

• the properties of the gas consisting of gas density (ρ) and gas viscosity (μ), and also

• the bed properties like bed voidage at minimum fluidization (εmf).

Many of the equations for calculating the minimum fluidization velocity are a function of two dimensionless numbers, Reynolds number (2.1) and Archimedes number (equation 2.2) [14, 15]

𝑅𝑒_𝑚𝑓= 𝜌𝑑_𝑝𝑈_𝑚𝑓

𝜇 2.1

𝐴𝑟 =𝜌𝑑_𝑝³(𝜌_𝑝− 𝜌)𝑔 𝜇²

2.2

Where 𝑑_𝑝 [m] is particle diameter, 𝜌_𝑝 [kg/m³] is particle density, 𝜇 [Pa.s] is the dynamic viscosity of air at ambient condition, 𝜌 [kg/m³] is the density of ambient air and 𝑔 [m/s²] is the gravitational acceleration.

One of the first equations to predict minimum fluidization velocity was developed by Ergun [16]. This equation is widely used for calculating the Umf and is based on the prediction of the pressure drop through a bed. This correlation is a semi-experimental equation to investigate the effects of the particle properties, gas characteristics, and bed properties on the minimum fluidization velocity, equation (2.3) [14].

1.75

𝜀_𝑚𝑓³ 𝜑𝑅𝑒_𝑚𝑓² +150(1 − 𝜀_𝑚𝑓)

𝜀_𝑚𝑓³ 𝜑² 𝑅𝑒_𝑚𝑓 = 𝐴𝑟 2.3

Where 𝜀_𝑚𝑓 [-] is bed voidage at minimum fluidization and 𝜑 [-] is sphericity.

Equation 2.3 can be expressed as equation 2.4 due to difficulties in finding the values of 𝜀_𝑚𝑓 and 𝜑 [14].

𝐾₁𝑅𝑒_𝑚𝑓² + 𝐾₂𝑅𝑒_𝑚𝑓 = 𝐴𝑟 2.4 Where K1 and K2 are equal to:

𝐾₁ = 1.75 𝜀_𝑚𝑓³ 𝜑

2.5

(24)

𝐾₂ =150(1 − 𝜀_𝑚𝑓) 𝜀_𝑚𝑓³ 𝜑²

2.6

One of the primary modified versions of the Ergun equation was presented by Wen and Yu [15]. They have found that for a wide range of Remf from 0.001 to 4000, the values of K1 and K2 remained unchanged. Therefore, equation (2.6) can be converted to equation (2.7).

𝑅𝑒_𝑚𝑓= (33.7²+ 0.0408𝐴𝑟)^0.5− 33.7 2.7

Where ^𝐾²

2𝐾₁ = 33.7 and ¹

𝐾₁ = 0.0408.

Hence, the general form of equation (2.7) can be like equation (2.8).

𝑅𝑒_𝑚𝑓 = (𝐾₁²+ 𝐾₂𝐴𝑟)^0.5− 𝐾₁ 2.8 Some equations in the form of equation (2.8) to predict the Umf are listed in Table 2.1.

Table 2.1: The correlations to predict minimum fluidization velocity [14]

Authors correlations

Particle diameter

𝑑_𝑝[𝜇𝑚]

Particle density 𝜌_𝑝[^𝑘𝑔

𝑠]

Geldart group

Bourgeis and Grenier [17] 𝑅𝑒𝑚𝑓= (25.46²+ 0.0382𝐴𝑟)^0.5− 25.46 86-25000 1200-19300 A, B, D Babu, Shah, and Talwalker [18] 𝑅𝑒𝑚𝑓= (25.25²+ 0.0651𝐴𝑟)^0.5− 25.25 50-2870 2560–3920 A, B, D Bin [19] 𝑅𝑒𝑚𝑓= (27.31²+ 0.0386𝐴𝑟)^0.5− 27.31 40-2120 1600-7500 A, B, D Hartman, Trnka and Svoboda

[20]

𝑅𝑒𝑚𝑓= (7.03²+ 0.0101𝐴𝑟)^0.5− 7.03 125-800 1700 A,B, D

Hilal, Ghannam and Anabtawi [21]

𝑅𝑒_𝑚𝑓= (13.07²+ 0.0263𝐴𝑟)^0.5− 13.07 80-1230 1228-8900 A,B,D

Saxena and Vogel [22] 𝑅𝑒_𝑚𝑓= (25.28²+ 0.0571𝐴𝑟)^0.5− 25.28 650-704 1900-2460 B, D Paudel and Feng [23] 𝑅𝑒_𝑚𝑓= (30.28²+ 0.0464𝐴𝑟)^0.5− 30.28 241-490 2500-3940 B

2.4.2 The experimental method to determine minimum fluidization velocity

The minimum fluidization velocity can be determined experimentally. The experimental measurement of Umf is simpler. At this point, the pressure drops through the particle bed or the

(25)

expanded bed height starts increasing significantly and experiences its maximum value, as shown in Figure 2.3. The pressure drop versus superficial gas velocity diagram is a useful indication of the state of fluidization and determining the minimum fluidization velocity [10].

Figure 2.3: Pressure drop vs superficial gas velocity [10]

2.5 Minimum bubbling velocity

When the gas velocity continues to increase, the bubbles start forming, as shown in Figure 2.1C, at this point the bubbling fluidized bed occurs at a superficial gas velocity very slightly higher than the minimum fluidization velocity. This kind of fluidization is called ‘aggregative fluidization’. In gas-solid bed with large size of particles, bubbles become visible at the same time with increasing gas velocity Umf, therefore for particles in Geldart group B and D, Umb is equal to Umf. In this type of fluidization, the bubble rising through the bed affects the movement of particles [8, 10].

Minimum bubbling velocity (Umb) is defined as a fluidizing velocity at which the first bubbles are observed. Abrahmsen and Geldart [21] did some experimental studies with 23 different powders with the diameter range of 23-76 µm and the densities of 1100-4600 kg/m³ and with 5 different gases. They proposed the correlation for minimum bubbling velocity with gas and particle properties, equation (2.9) [24, 25].

𝑈_𝑚𝑏= 2.07 exp(0.716 𝐹₄₅)𝑑̅_𝑝𝜌_𝑔^0.06

𝜇^0.347 2.9

Where, 𝐹₄₅ is the mass fraction of particles having the diameter of particles smaller than 45 µm.

(26)

2.6 Terminal settling velocity

The velocity of a particle in a steady, non-fluctuating fluid flow will reach a constant value that it is the maximum velocity associated with a particle or droplet fall into a fluid. When particles move in a stagnant medium (uF=0, the fluid velocity), at some speed, the drag or force of resistance will equal the gravitational pull on the object [26].

The drag force FD on a particle fall through a fluid is given by equation (2.10).

𝐹_𝐷 = 1

2𝐶_𝐷𝜌_𝑔𝑎𝑠𝑣²𝐴_{𝑝,𝑝𝑟𝑜𝑗} 2.10

Where,

𝐶_𝐷[-] = Drag coefficient 𝑣 [m/s] = settling velocity

𝐴_{𝑝,𝑝𝑟𝑜𝑗} [m²] = Projected area of the particle

By assuming spherical particles and settling in a Stokes regime, where relatively small particles are moving in a fluid and the Reynolds number is low, Re<<1, the settling velocity can be calculated by equation (2.11).

𝑣_𝑡 =𝑔 𝑑_𝑝²(𝜌_𝑝− 𝜌_𝑔𝑎𝑠)

18𝜇 2.11

For the bigger particles (Re>>1), the settling is turbulent and terminal velocity can be calculated by applying equations (2.12) and (2.13) [26].

𝑅𝑒 = 0.1334 𝐴𝑟^0.7016 2.12

𝑣_𝑡 =𝑅𝑒𝜇 𝜌𝑑_𝑝

2.13

2.7 The impact of high temperature on the fluidization

At high temperatures, the agglomeration and sintering of particles will occur and the behavior of the fluidization can be changed. This phenomenon should be considered in industrial applications especially coal ash and metal with a wide variety of impurities. The sintering of particles can lead to undesirable behavior due to forming low-melting eutectics of materials at the surface of particles. Despite the disadvantages of sintering, it is applied to develop processes for the agglomeration of fine particles [10].

In addition, a higher temperature will change the viscosity and density of the gas, and this will change the fluidization characteristics of the powders, for example via the impact on Re and Ar numbers according to equations 2.1 and 2.2.

(27)

2.8 Fluidization of Geldart C particles

Geldart C particles are difficult to fluidize and most time, they tend to grow as a plug of particles. When the diameter of beds is large, and the particles are exposed to the fluidized gas, cracks, channels, or rat holes are formed with no fluidization of particles. The fluidization behavior of C powders is due to the vigorous inter-particle forces. When the surface-to-volume ratio of the particles increases, the interparticle forces get larger, and consequently the distance between particles reduces. The cohesive forces of C powders become greater than the particle gravity and the hydrodynamic forces applied by fluidization gas around the particle [10, 27].

Different approaches can be introduced to fluidize fine particles (Geldart C), such as:

1. According to Kunii and Levenspiel [10], “One way of processing these solids is to introduce them into a bed of the same material but of larger size, preferably Geldart B. Even though the fines are very small, they are not entrained immediately but may stay in the bed an average of several minutes. This usually is long enough for a physical or chemical transformation of these solids.”

2. As a result of the Kunii’s approach, Kato et al [28] have proposed a new type of fluidized bed, “Powder-Particle Fluidized Bed, PPFB”, in order to fluidize fine powders of group C. They used Gelart B particles (inert particles) to provide a homogeneous distribution of the fine particles and enhance the flowability. The aggregation and clustering of the fine particles will decrease due to collisions with inert coarse particles [28].

3. Tashimo et al [29] have been applied PPFB to calcine the small limestone particles in the range of 2-64 µm. The effects of reaction temperature, size of limestone powders, superficial gas velocity, and bed height on the calcination conversion were studied. The sea sand particle (inert particles) with the diameter of 420-840 µm, minimum fluidization velocity Umf=0.14 m/s, and terminal velocity Ut= 5.85 m/s, the static bed height is in the range of 0.10 to 0.20 m, and T=1073 K were applied to enhance the flowability of the fine powders. The mean sizes of the powders have consisted of 2.0, 5.0, 9.9, 23.3, 30.9, and 64 µm and the component of CaCO3 in the limestone was assumed 99%. Figure 2.4 depicts the schematic of the experimental apparatus which has been used in the paper. As can be seen in the figure, the system consists of a powder-feeder, a powder-particle fluidized bed reactor heated by an electric furnace controlled by a P.I.D controller, and a filter for separating the product powder of the calcined meal. The internal diameter of the reactor is equal to 0.03 m and the height of the stainless-steel column is 0.65 m. At the bottom, there is a porous distributor with 10 µm holes.

The procedure of the experiment was that a combination of air and CO2 was blown to the column, the coarse particles were fluidized with the gas. At the steady-state condition, the limestone powder was fed pneumatically into the bed, 0.02 m above the distributor. The superficial gas velocities varied in the range of 0.25 to 1.0 m/s.

the experimental conditions are represented in Table 2.2. The CaO powder entrained from the top and collected by the bag filter and outside of the reactor before entering bag filters was cooled with the water.

(28)

Table 2.2: Experimental conditions of tests done by Tashimo et al [29]

Operating factors Ranges

Reaction temperature, T [K] 1073-1223

Mean diameter of limestone, dp[µm] 2-64 Superficial gas velocity, U [m/s] 0.25 – 1

Static bed height, L [m] 0.1 – 0.2

Feeding rate, Fr [g/h] 5-15

Size of inert particles [µm] 420-840

Figure 2.4: Schematic diagram of the experimental apparatus [29]

4. Another approach to granulate fine powders could be based on the mechanical agitation concept. One practical way is to use chopper in the fluidized bed to benefit from the high shear mixer with the fluidized beds’ advantages [27, 30].

5. One common way in the pharmaceutical industry is using Vibro-fluidization. In this case, a rotary fluidized bed granulator is applied to transfer vibration energy to the bed. The drawback of this system is that the fluidized bed is inhomogeneous and the temperature gradient is observed in the bed [27, 31].

6. One approach to handle the cohesive forces of the fine particle can be to reduce the granular Bond number. The granular Bond number is defined as the ratio of the interparticle attractive force to particle weight or particle body force to enhance the

(29)

flowability. Two methods have been proposed by the New Jersey Institute of Technology group [32, 33] to decrease the granular Bond number; 1) applying a centrifugal field to increase the particle body force and 2) dry particle coating in order to handle inter-particle cohesion. Chen et al [27] have introduced a new technique to fluidize very fine particles with a size of 15 µm. the nanosized particles have been applied to pre-coat the fine particles to be granulated. Nanosized particles can be covered the surface of the fine particles and decrease the cohesive forces [27].

2.9 Heat transfer in fluidized beds

The high heat and mass transfer rate and good mixing properties are the main merits of the fluidized beds. There are several different processes regarding heat transfer in a fluidized bed, such as particle-gas heat transfer, heat transfer between the solid particles and larger particles floating in the bed, heat transfer between different points in the bed, and transfer of heat between submerged surfaces and the bed [34, 35].

In the high-temperature uniform fluidization, there is a low-temperature difference (2-5 °C) between different points in the bed. Due to the high capacity of the solid particles to transfer the heat to the fluidizing gas, the gas temperature becomes equal to particle temperature when leaving the bed. Consequently, the large heat capacity of the particles leads to a very small temperature difference between gas and solid. In order to calculate the gas to particle heat transfer coefficient, the Gelperin and Einstein relation can be applied, equations (2.14) and (2.15) [35].

𝑁𝑢_𝑝 = 1.6 × 10⁻²(𝑅𝑒_𝑝 𝜀 )

1.3

𝑃𝑟^0.33 for 𝑅𝑒_𝑝

𝜀 < 200 2.14

𝑁𝑢_𝑝 = 0.4 (𝑅𝑒_𝑝 𝜀 )

2

3 𝑃𝑟^0.33 for 𝑅𝑒_𝑝

𝜀 > 200 2.15 Where Nup [-] is Nusselt number and Pr [-] is Prandtl’s number and is defined as equation (2.16).

𝑃𝑟 =𝜇 𝑐_𝑝

𝑘 2.16

That 𝜇 [kg/ms] is dynamic viscosity, 𝑐_𝑝 [J/kgK] is specific heat and 𝑘 [W/mK] is thermal conductivity.

2.9.1 Heat transfer in bubbling fluidized beds

In most industrial applications of the bubbling fluidized beds, numerous heating tubes are utilized to heat transfer in the beds. The orientation of the heating tubes can be horizontal or vertical. In [36], some methods to predict the heat transfer coefficient on submerged tubes have

(30)

been proposed. It is concluded that heat transfer occurs by gaseous convection through contact with a fluidized gas (lean gas phase), by particle conduction or convection during the times in contact with the dense phase of solid, and radiation heat transfer at high-temperature. Hence, the effective heat transfer coefficient (ℎ_𝑡𝑜𝑡) can be affected by the heat transfer coefficient during “dense” (particle) phase (ℎ_𝑑) contact, heat transfer coefficient during “lean” gas-phase contact (ℎ_𝑙), and heat transfer coefficient for radiation (ℎ_𝑟). Equation (2.17) represents the correlation for the effective heat transfer coefficient [36].

ℎ_𝑡𝑜𝑡= 𝑓_𝑑ℎ_𝑑 + (1 − 𝑓_𝑑)ℎ_𝑙+ ℎ_𝑟 2.17

Where 𝑓_𝑑 is the time fraction of contact by dense phase and it is assumed that hat the three heat transfer coefficients on the right-hand side above are independent.

2.10 The regular cement kiln system

Figure 2.5 illustrates a process flow diagram of a regular cement kiln system. Two preheater towers by using hot exhaust gas of calciner can heat the raw meal from approximately 50 to 700 °C. In the calciner, the energy of fuel combustion supplies the required heat for the endothermic calcination (decarbonation) process (CaCO3→CaO+CO2) at a temperature of 900

°C. Therefore, the CO2 is produced during two stages; fuel combustion and calcination reaction. The calcination process is completed in the rotary kiln where precalcined meal with the calcination degree of 90-95% enters and 100% calcination happens a few meters into the rotary kiln, after that the temperature of the meal starts to rise, and melt-phase is formed.

Eventually, clinker is formed. Due to high-temperature conditions, some of the meal melts and form the clinker nodules inside the kiln. The temperature of the gas can reach up to 2000 °C.

In the rotary kiln, the gas and solids flow in the opposite direction to another, and when the solids enter the kiln some bypass gas exit the system. In the cooler, by applying the ambient air, the temperature of the clinker can decrease from 1400 °C to 100 °C. Some part of the heater air in this section is used as secondary air in the rotary kiln and also as tertiary air in the calciner and still some part of cooling air is evacuated to the environment [37].

(31)

Figure 2.5: A cement regular kiln process with two preheater strings [37]

2.11 Generic electrified calciner system

Figure 2.6 depicts a modified cement kiln process where the electrical energy is applied for decarbonating the raw meal in the calciner. This project is a part of Phase 2 of the ELSE¹ project associated with the collaboration of the University of South-Eastern Norway and SINTEF Tel-Tek in the project of utilizing green electricity instead of carbon-containing fuels.

By applying electricity not only the CO2 production from the fuel combustion can be eliminated but also provide a basis for simple capture of the CO2 generated through calcination, as CO2 is the only gaseous product exiting from the electrified calciner [6, 37].

1 “ELSE is short for ‘ELektrifisert SEmentproduksjon’ (Norwegian) meaning ‘electrified cement production’ ”.

(32)

Figure 2.6: A modified cement kiln process applying electrical energy for calcination [37]

Calcination in an electrically heated bubbling fluidized bed applied in calcium looping

Faculty of Technology, Natural sciences and Maritime Sciences

Campus Porsgrunn FMH606 Master's Thesis 2020 Energy and Environmental Technology