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Characterization of in-line mixing of pulp fibre suspensions based on electrical resistance tomography Yenjaichon, Wisarn 2012

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CHARACTERIZATION OF IN-LINE MIXING OF PULP FIBRE SUSPENSIONS BASED ON ELECTRICAL RESISTANCE TOMOGRAPHY by  Wisarn Yenjaichon B.Eng., Chulalongkorn University, 2002  A THESIS SUBMITTED IN PARTIAL FULFILMENT OF THE REQUIREMENTS FOR THE DEGREE OF DOCTOR OF PHILOSOPHY in The Faculty of Graduate Studies (Chemical and Biological Engineering)  THE UNIVERSITY OF BRITISH COLUMBIA (Vancouver) December 2012  © Wisarn Yenjaichon, 2012  Abstract In pulp bleaching processes, pre-distribution of chemicals in suspensions ahead of tower reactors is essential to ensure efficient lignin removal and optimal use of the chemicals. In-line mixers, combined with chemical injectors, are commonly used to achieve this goal. In spite of its importance, in-line mixing of pulp suspensions is not well understood. In this thesis, liquid distribution and gas dispersion were investigated downstream of in-line mixers, including jet and mechanical mixers, to provide better understanding and guidance for mixer design and process optimization. In the present work, non-intrusive electrical resistance tomography (ERT) was used to quantify mixing based on two novel mixing indices, derived from the standard deviation of image pixel values. This technique was also implemented as a real-time mixing assessment tool in industrial pulp bleaching, with success in monitoring mixing quality as a function of process operating conditions. Liquid jet mixing was found to depend strongly on the flow regime and jet penetration. For turbulent flow, the criteria for in-line jet mixing in water apply also to suspensions. When a suspension flows as a plug, mixing differs greatly from that in water, depending on the fibre network strength in the core of the pipe. With an impeller present, mixing improved substantially, primarily in the high-shear zone around the impeller, with rapid reflocculation downstream. Gas mixing depended on the flow regime and buoyancy in a complex manner. When buoyancy was not significant, impeller operation enhanced mixing  ii  since bubbles dispersed throughout the pipe cross-section, whereas without the impeller, the bubbles congregated near the wall due to robust fibre networks in the core of the pipe. For buoyancy-dominated flow, the impeller worsened mixing since it disrupted the fibre networks and delivered gas to the top of the pipe, whereas the networks caused liquid/pulp slugs to flow at the top for a tee alone.  iii  Preface Part of Chapter 1, together with Chapter 8, was published as Yenjaichon, W., Pageau G., Bhole, M., Bennington, C.P.J, Grace, J.R, 2011. Assessment of mixing quality for an industrial pulp mixer using electrical resistance tomography. Can. J. Chem. Eng. 89 (5), 996–1004. The sensor system was designed by Chad Bennington and me, and was fabricated and installed in the first chlorine dioxide bleaching (D0) stage at Howe Sound Pulp and Paper Ltd. by Gerry Pageau. The experiments were conducted by Gerry Pageau, Manish Bhole and me, and data analysis was performed by me, with guidance from Chad Bennington and Manish Bhole. I prepared the manuscript, with insightful feedback from John Grace, Manish Bhole and Gerry Pageau. Parts of sections 3.1 and 3.3.1 and parts of Chapter 4 were published: Yenjaichon, W., Grace, J.R., Lim, C.J., Bennington, C.P.J., 2012. In-line jet mixing of liquid-pulp-fibre suspensions: Effect of concentration and velocities. Chem. Eng. Sci. 75, 167–176. The test section and electrical resistance tomography system were designed by me, with helpful advice from Chad Bennington. I was responsible for data collection, analysis and manuscript preparation. John Grace and Jim Lim provided feedback and guidance throughout this process. Parts of Chapter 4 have been accepted for publication: Yenjaichon, W., Grace, J.R., Lim, C.J., Bennington, C.P.J., 2012. In-line jet mixing of liquid-pulpfiber suspensions: Effect of fiber properties, flow regime and jet penetration, AIChE J. Chad Bennington provided advice on the design of experimental equipment. All  iv  the experiments and data analysis were conducted by me, with guidance from John Grace and Jim Lim. The written work was a collaborative effort by John Grace, Jim Lim and me. Part of section 3.1 and parts of Chapter 5 have been accepted for publication: Yenjaichon, W., Grace, J.R., Lim, C.J., Bennington, C.P.J., 2012. Pilotscale examination of mixing liquid into pulp fiber suspensions in the presence of an in-line mechanical mixer, Ind. Eng. Chem. Res. The mechanical mixer system was designed by the author, with guidance from Profs. Bennington, Grace and Lim. I was responsible for collecting and analyzing data, and preparing the manuscript. John Grace and Jim Lim revised the manuscript and provided insightful feedback. Part of section 3.3.2 and parts of Chapter 6 have been accepted for publication: Yenjaichon, W., Grace, J.R., Lim, C.J., Bennington, C.P.J., 2012. Characterisation of gas mixing in water and pulp-suspension flow based on electrical resistance tomography, Chem. Eng. J. The gas injection system was designed by the author, with helpful advice from Chad Bennington. I was responsible for data collection, analysis and manuscript preparation. John Grace and Jim Lim provided feedback throughout this process. Parts of Chapter 6 have been accepted for publication: Yenjaichon, W., Grace, J.R., Lim, C.J., Bennington, C.P.J., 2012. Gas dispersion in horizontal pulpfibre-suspension flow, Int. J. Multiphase Flow. Chad Bennington provided guidance on the design of experimental equipment. I conducted all the experiments and data analysis. The written work is a collaborative effort by John Grace, Jim Lim and me. Parts of Chapter 7 are based on a paper submitted for publication.  v  Table of Contents Abstract................................................................................................................... ii Preface.................................................................................................................... iv Table of Contents................................................................................................... vi List of Tables......................................................................................................... xii List of Figures....................................................................................................... xiii Nomenclature................................................................................................. xxxviii Acknowledgements.............................................................................................. xlii 1. Introduction………………………………………………………………………….. 1 1.1 Thesis Objectives…………………………………………………………………. 3 1.2 Thesis Organization………………………………………………………………. 4 2. Background and Literature Review……………………………………………... 8 2.1 Pulp Bleaching Processes.………………………………………………………. 8 2.2 Mixing Equipment in Pulp Bleaching Processes……………………………... 10 2.2.1  Low-consistency mixers………………………………………………… 11  2.2.2  Medium-consistency mixers….……………………………………….... 12  2.2.3  High-consistency mixers.……………………………………………….. 13  2.3 Benefits of Improved Mixing……………………………………………………. 13 2.4 Mixing Scales…………………………………………………………………….. 14 2.5 Measurement of Mixing Quality………………………………………………… 15 2.6 Mixing Assessment on Laboratory and Pilot Scales of Pulp Suspensions....17 2.7 Mixing Assessment on Industrial Scale in Pulp and Paper Processes……. 19  vi  2.8 Electrical Resistance Tomography (ERT)……………………………............. 21 2.8.1  Principles and system structure……………………………….……….. 21  2.8.2  ERT applications on the industrial scale………………………………. 26  2.8.3  Use of ERT for quantification of mixing……………………………….. 27  2.8.4  Use of ERT in pulp and paper systems……………………………….. 30  3. Experimental Details and Mixing Indices……………………………………... 34 3.1 Experimental Apparatus………………………………………………………… 34 3.2 Materials and Experimental Methods………………………………………….. 42 3.2.1  Test procedures for liquid injection…………………………………….. 43  3.2.2  Test procedures for gas injection…………………………………….... 44  3.3 Data Analysis…………………………………………………………………….. 45 3.3.1  Characterization of liquid-suspension mixing………….……………... 47  3.3.2  Characterization of gas-suspension mixing…………………………... 49  4. In-Line Jet Mixing of Liquid-Pulp-Fibre Suspensions………………………. 53 4.1 Introduction……………………………………………………………………….. 53 4.2 Experimental Details…………………………………………………………….. 60 4.3 Results and Discussion………………………………………………………..... 60 4.3.1  Jet mixing in Newtonian fluid (water)……… ………….……………..,. 60  4.3.2  Flow regime of pulp suspension flow in the pipe…………………...... 69  4.3.3  Effect of mainstream velocity on mixing quality in softwood pulp suspensions………………………………………………………………. 73  4.3.4  Effect of jet velocity on mixing in softwood pulp suspensions………. 79  vii  4.3.5  Effect of velocities on mixing quality for hardwood pulp suspensions and optimum operating conditions for both pulp types………………. 83  4.3.6  Effect of fibre mass concentration on mixing quality……………….… 89  4.3.7  Effect of fibre-turbulence interactions on mixing quality for hardwood pulp………………………………………………………………………… 92  4.3.8  Effect of fibre type on mixing quality…………………………………… 99  4.4 Conclusions…………………………………………………………………….. 100 5. Mixing Liquid into Pulp Fibre Suspensions in the Presence of an In-Line Mechanical Mixer………………………………………………………………… 104 5.1 Introduction……………………………………………………………………… 104 5.2 Experimental Details…………………………………………………………… 106 5.3 Results and Discussion………………………………………………………... 106 5.3.1  Effect of velocities and impeller rotational speed on mixing quality in water.,……………………………………………………………………. 106  5.3.2  Effect of mainstream velocity on mixing in pulp suspensions.......... 113  5.3.3  Effect of jet velocity on mixing quality in pulp suspensions.............. 118  5.3.4  Effect of impeller rotational speed on mixing quality in pulp suspensions…………………………………………………………….. 121  5.3.5  Effect of fibre mass concentration on mixing quality……………….. 128  5.3.6  Hardwood vs. softwood fibres………………………………………… 133  5.4 Conclusions…………………………………………………………………….. 135 6. Gas Dispersion in Horizontal Pulp-Fibre-Suspension Flow……………… 139 6.1 Introduction……………………………………………………………………… 139  viii  6.2 Experimental Details…………………………………………………………… 144 6.3 Results and Discussion………………………………………………………... 145 6.3.1  ERT for evaluating gas-liquid horizontal flow...,………….…………. 145  6.3.2  Gas mixing indices……………………………………………….......... 147  6.3.3  Effect of injection tube geometry…………………………….............. 158  6.3.4  Effects of salt and surfactant on bubble coalescence...................... 159  6.3.5  Overall gas holdup……………………………………………………… 160  6.3.6  Effect of fibre mass concentration on gas mixing in pulp suspension flow……………………………………………………………………….. 163  6.3.7  Flow regime for air-suspension flow and effect of superficial liquid/pulp velocity…………………………………………………………………… 172  6.3.8  Effect of superficial gas velocity on mixing in pulp suspensions….. 177  6.4 Conclusions…………………………………………………………………….. 180 7. Gas Dispersion in Pulp-Suspension flow in the Presence of an In-Line Mechanical Mixer………………………………………………………………… 182 7.1 Introduction……………………………………………………………………… 182 7.2 Experimental Details…………………………………………………………… 183 7.3 Results and Discussion………………………………………………………... 184 7.3.1  Effect of velocities and rotational speed on gas mixing in water flow...,……………………………………………………………………. 184  7.3.2  Effect of fibre mass concentration on gas mixing in pulp suspensions……………………………………………………............. 188  7.3.3  Effect of superficial gas velocity on gas-suspension mixing............ 194  ix  7.3.4  Influence of superficial liquid/pulp velocity…………………………… 202  7.3.5  Effect of impeller speed on gas-suspension mixing………………… 204  7.4 Conclusions…………………………………………………………………….. 209 8. Mixing Quality for an Industrial Pulp Mixer using Electrical Resistance Tomography………………………………………………………………………. 212 8.1 Introduction……………………………………………………………………… 212 8.2 Experimental Details…………………………………………………………… 215 8.3 Results and Discussion………………………………………………………... 218 8.3.1  New mixer installation at HSPP………………………………………. 218  8.3.2  Reaction of ClO2 with pulp……………………………………............. 221  8.3.3  Variability of conductivity with time for pulp suspension and ClO2 feed to mixer............................................................................................ 223  8.3.4  Flow regime……………………………………………………………... 224  8.3.5  Mixing quality as a function of process operating conditions........... 224  8.4 Conclusions…………………………………………………………………….. 232 9. Conclusions, Contributions and Recommendations……………………… 234 9.1 Conclusions…………………………………………………………………….. 234 9.1.1  Electrical resistance tomography as a measurement technique….. 234  9.1.2  Liquid jet mixing into liquid mainstream.………………..................... 235  9.1.3  Gas mixing into liquid mainstream................................................... 237  9.2 Contributions and Potential Applications…………………………………….. 240 9.3 Recommendations for Future Research…………………………………….. 241 References........................................................................................................... 244  x  Appendix A: Supplementary Data for Chapter 4.............................................. 262 Appendix B: Supplementary Data for Chapter 5………………...……………… 271 Appendix C: Supplementary Data for Chapter 6............……………………….. 282 Appendix D: Supplementary Data for Chapter 7..........................................… 290  xi  List of Tables Table 2.1: Typical operating conditions of first chlorine dioxide (D0) and chlorine dioxide D1 bleaching stages (liquid-suspension applications) at Howe Sound Pulp and Paper (HSPP) Ltd. (Data courtesy of the company)..... 9 Table 2.2: Typical operating conditions of oxygen delignification (O) and oxidative alkaline extraction (EO) stages (gas-suspension applications) at HSSP (Data courtesy of the company)............................................................. 10 Table 4.1: Summary of previous studies on in-line jet mixing for Newtonian fluids (updated from Forney, 1986).................................................................. 56 Table 4.2: Velocity ratios, R, for optimum mixing at Up = 0.5 m/s, Dr = 0.05........... 88 Table 4.3: Velocity ratios for jet reaching centre of pipe for Dr = 0.05..................... 90 Table 5.1: Summary of the influences of jet penetration, flow regime and impeller speed on mixing …………………………………………………………... 136 Table 6.1: Summary of techniques for measuring void fraction and bubble size for different flow patterns in air-water horizontal flow (DB = dispersed bubble, SL = slug, S = stratified, W = wavy)........................................ 142 Table 6.2: Comparison of gas holdup measured without resistor adaptors, with adaptors and visual observation for the stratified flow pattern..............147 Table 6.3: Flow regimes (identified as in Figure 6.6) for air-suspension horizontal flow for Usg = 0.11 and 0.44 m/s, x/D = 22.1 and various fibre mass concentrations.......................................................................................173 Table 8.1: Comparison of mixing assessment of industrial pulp mixers…………. 231  xii  List of Figures Figure 2.1: Process diagram of a typical bleaching stage (adapted from Bennington, 2004)...................................................................................................... 9 Figure 3.1: Schematic of pilot-scale flow loop facility.............................................. 35 Figure 3.2: Schematic of test section (all dimensions in mm)................................. 36 Figure 3.3: Photo showing test section and ERT system........................................ 37 Figure 3.4: Parameter setup for electrical resistance tomography.......................... 38 Figure 3.5: Schematic of in-line mechanical mixer (all dimensions in mm)............. 41 Figure 3.6: Electrical conductivity as a function of NaCl concentration................... 44 Figure 3.7: ERT image reconstruction grid for pipe circular cross-section (ITS, 2007)………………………………………………………………………... 46 Figure 3.8: Tomographic image showing regions of high and low conductivity (ITS, 2007). Bar at the bottom shows colours corresponding to different conductivities........................................................................................ 46 Figure 4.1: Modified mixing index as a function of dimensionless distance downstream, x/D, for Newtonian fluid (water) at Up = 1.0 m/s, 2.0 m/s and 3.0 m/s for virtually identical jet-to-pipe velocity ratio and Dr = 0.05. Locations of planes P2 to P8 are shown in Figure 3.2......................... 61 Figure 4.2: Tomographic images for Newtonian fluid (water) for Dr = 0.05 at (a) Up = 1.0 m/s, R = 4.50, (b) Up = 2.0 m/s, R = 4.41 and (c) Up = 3.0 m/s, R = 4.23. Locations of planes P1 to P8 are shown in Figure 3.2................ 62  xiii  Figure 4.3: Modified mixing index as a function of dimensionless distance downstream of injection for various jet-to-pipe velocity ratios, R, for water flow at Dr = 0.05.......................................................................... 64 Figure 4.4: Tomographic images for (a) near-wall (wall-source) injection with R = 2.16, (b) jet-mixing with R = 6.39 and (c) jet-impaction with R = 15.5, Dr = 0.05 in water flow.............................................................................. 65 Figure 4.5: Jet-to-pipe velocity ratios, R, as a function of jet-to-pipe diameter ratio for the jet reaching the axis of the pipe for 90-tee mixers................... 66 Figure 4.6: Comparison of long and short injection tubes on modified mixing index at Up = 1.0 m/s, Uj = 2.0 m/s, Dr = 0.167 for fibre-free water............... 67 Figure 4.7: Tomographic images for water flow at Up = 1.0 m/s, Uj = 2.0 m/s, Dr = 0.167: (a) short injection tube; (b) long injection tube.......................... 68 Figure 4.8: Typical head loss-velocity curve for water and low-consistency pulp suspension flow through pipes (adapted from Robertson and Mason, 1957). Dashed line is the corresponding line for water........................ 71 Figure 4.9: Head loss-velocity curve for water and softwood pulp suspensions in this work...................................................................................................... 72 Figure 4.10: Head loss-velocity curves for water and hardwood kraft pulp............. 73 Figure 4.11: Modified mixing index as a function of dimensionless distance downstream of injection for softwood pulp suspension at Cm = 0.5% and for water, Dr = 0.05 with various mainstream velocities and similar jet-topipe velocity ratios................................................................................ 74 Figure 4.12: Tomographic images for softwood pulp suspension flow with Cm = 0.5%, Dr = 0.05 at the following: (a) Up = 1.0 m/s, R = 3.63; (b) Up = 2.0 m/s, R = 3.36; (c) Up = 4.0 m/s, R = 3.16; and (d) Newtonian fluid  xiv  (water) at Up = 4.0 m/s, R = 3.17. Locations of planes P1, P2, P5 and P8 are shown in Figure 3.2.................................................................. 76 Figure 4.13: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension with Cm = 2.0% and for water, Dr = 0.05 for various mainstream velocities and similar jet-to-pipe velocity ratios....................................................................................... 77 Figure 4.14: Tomographic images for softwood pulp suspension flow with Cm = 2.0%, Dr = 0.05 at the following: (a) Up = 1.0 m/s, R = 3.77; (b) Up = 2.0 m/s, R = 3.36; (c) Up = 4.0 m/s, R = 3.22; and (d) Newtonian fluid (water) at Up = 4.0 m/s, R = 3.17. Locations of planes P1, P2, P5 and P8 are shown in Figure 3.2.................................................................. 78 Figure 4.15: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Cm = 0.5% and for water, Up = 2.0 m/s, Dr = 0.05 for various jet velocities in wall-source and jetmixing modes....................................................................................... 80 Figure 4.16: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Cm = 0.5% and for water, Up = 0.5 m/s, Dr = 0.05 for various jet velocities in jet-mixing and jetimpaction modes.................................................................................. 80 Figure 4.17: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Cm = 2.0%, Up = 0.5 m/s, Dr = 0.05 for various jet velocities in jet-mixing and jet-impaction mode.................................................................................................... 81 Figure 4.18: Tomographic images for softwood pulp suspension flow at Up = 0.50 m/s for Cm = 2.0%, Dr = 0.05 at the following: (a) R = 9.62; (b) R = 13.5; (c) R = 16.9; and (d) R = 24.9. Locations of planes P1, P2, P5 and P8 are shown in Figure 3.2........................................................................ 82  xv  Figure 4.19: Modified mixing index as a function of dimensionless distance downstream at different fibre mass concentrations of hardwood pulp suspensions: (a) Cm = 1.0% and (b) Cm = 3.0% for Dr = 0.05, almost identical jet-to-pipe velocity ratios and various mainstream velocities, and comparison with modified mixing index for water under very similar experimental conditions....................................................................... 84 Figure 4.20: Modified mixing index as a function of dimensionless distance downstream at different fibre mass concentrations of hardwood pulp suspensions: (a) Cm = 1.0% and 0% and (b) Cm = 3.0% for Up = 0.5 m/s, Dr = 0.05 and various jet velocities.............................................. 86 Figure 4.21: Tomographic images for hardwood pulp suspension flow at Up = 0.50 m/s for Cm = 3.0%, Dr = 0.05 at: (a) R = 8.2; (b) R = 12.3; (c) R = 16.0; and (d) R = 24.1. The locations of planes P1, P2, P5 and P8 are shown in Figure 3.2......................................................................................... 87 Figure 4.22: Modified mixing index as a function of dimensionless distance downstream at Up = 4.0 m/s, R ≈ 3.2 for various fibre mass concentrations...................................................................................... 91 Figure 4.23: Modified mixing index as a function of dimensionless distance downstream for Up = 1.0 m/s, virtually identical velocity ratios and various fibre mass concentrations........................................................ 91 Figure 4.24: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension at Cm = 0.5%, R = 3.1 compared with water for various mainstream velocities and almost identical jet-to-pipe velocity ratios........................................................ 92 Figure 4.25: Modified mixing index as a function of dimensionless distance downstream for water (W) and hardwood pulp suspension (HW) flow at Cm = 0.5%, Up = 3.0 m/s with various jet velocities.............................. 93  xvi  Figure 4.26: Modified mixing index as a function of dimensionless distance downstream for water, softwood and hardwood pulp suspensions at Cm=0.5%, Up = 4.0 m/s, R = 3.1 and Dr = 0.05.................................... 94 Figure 4.27: Tomographic images for Cm=0.5%, Up = 4.0 m/s, R = 3.1 and Dr = 0.05 with: (a) softwood; (b) hardwood; and (c) fibre-free water. The locations of planes P2, P4, P6 and P8 are shown in Figure 3.2......................... 97 Figure 4.28: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension at Up = 4.0 m/s, R = 3.1 with various fibre mass concentrations........................................................ 99 Figure 4.29: Modified mixing index as a function of dimensionless distance downstream for water, softwood and hardwood pulp suspensions at (a) Cm = 1.0% and (b) Cm = 3.0%, for Up = 3.0 m/s and R = 3.1.............. 101 Figure 5.1: Modified mixing index as a function of dimensionless distance downstream for Newtonian fluid (water) at Up = 1.0, 2.0 and 3.0 m/s, Dr = 0.05 with: (a) perpendicular static impeller (b) impeller rotating at N = 400 rpm. Locations of planes P2 to P8 are shown in Figure 3.5........ 107 Figure 5.2: Modified mixing index as a function of dimensionless distance downstream for water at Up = 1.0, 2.0 and 3.0 m/s, Dr = 0.05, virtually identical velocity ratios with: (a) N = 600 rpm (b) N = 800 rpm.......... 109 Figure 5.3: Modified mixing index as a function of dimensionless distance downstream of injection for water at Up = 1.0 m/s, Dr = 0.05, almost constant velocity ratios and various rotation speeds.......................... 110 Figure 5.4: Modified mixing index as a function of dimensionless distance downstream of injection for water, Dr = 0.05, almost constant velocity ratios and various rotation speeds at: (a) Up = 2.0 m/s (b) Up = 3.0 m/s......................................................................................................111  xvii  Figure 5.5: Modified mixing index as a function of dimensionless distance downstream of injection for various jet-to-pipe velocity ratios, R, with water at Dr = .05, N = 400 rpm........................................................... 113 Figure 5.6: Tomographic images for water, Dr = 0.05, N = 400 rpm at: (a) R = 6.15 (b) R = 12.2 (c) R = 24.6. Locations of planes P1 to P8 are shown in Figure 3.5........................................................................................... 114 Figure 5.7: Modified mixing index as a function of dimensionless distance downstream of injection for softwood pulp suspension with Cm = 0.5% and water (w), Dr = 0.05, N = 400 rpm with various mainstream velocities and almost identical jet-to-pipe velocity ratios.................... 116 Figure 5.8: Modified mixing index as a function of dimensionless distance downstream with and without impeller for softwood pulp suspension for Cm = 2.0% and water (w), Dr = 0.05, various mainstream velocities and almost identical jet-to-pipe velocity ratios and rotational speeds....... 117 Figure 5.9: Modified mixing index as a function of dimensionless distance downstream of injection for softwood pulp suspension with Cm = 0.5%, Up = 1.0 m/s, Dr = 0.05 and various jet velocities............................... 118 Figure 5.10: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension with Cm = 0.5% and water (w), Up = 0.5 m/s, Dr = 0.05 and various jet velocities........................ 119 Figure 5.11: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Cm = 2.0%, Up = 0.5 m/s, Dr = 0.05, almost constant impeller rotation speeds and various jet velocities............................................................................................ 121 Figure 5.12: Tomographic images for softwood pulp suspension at Up = 0.5 m/s, Cm = 2.0%, Dr = 0.05 with (a) R = 12.2, N = 403 rpm and (b) R = 16.5, N = 410 rpm. Locations of planes P1 to P8 are shown in Figure 3.5…… 122 xviii  Figure 5.13: Modified mixing index as a function of dimensionless distance downstream of injection for hardwood pulp suspension with Cm = 1.0% and water (w), Dr = 0.05 and almost constant velocity ratios with various impeller speeds at: (a) Up = 0.5 m/s (b) Up = 2.0 m/s..........................124 Figure 5.14: Modified mixing index as a function of dimensionless distance downstream of injection for softwood pulp suspension with Cm = 3.0% and water (w), Dr = 0.05 and almost identical velocity ratios with various impeller speeds at: (a) Up = 1.0 m/s (b) Up = 2.0 m/s......................... 126 Figure 5.15: Modified mixing index as a function of dimensionless distance downstream of injection for softwood pulp suspension with Cm = 3.0%, Dr = 0.05 and similar velocity ratios for various impeller speeds and mainstream velocities......................................................................... 127 Figure 5.16: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Up = 1.0 m/s, R = 6, N = 400 rpm, Dr = 0.05 for various fibre mass concentrations.................. 128 Figure 5.17: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension at Up = 1.0 m/s, R = 6, N = 800 rpm, Dr = 0.05 for various fibre mass concentrations.................. 129 Figure 5.18: Tomographic images for hardwood pulp suspension at Up = 1.0 m/s, R = 6, N = 800 rpm, Dr = 0.05 with (a) Cm = 0.5% and (b) Cm = 3.0%. Locations of planes P1 to P8 are shown in Figure 3.5………………. 131 Figure 5.19: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Up = 3.0 m/s, R = 4, Dr = 0.05 and virtually identical rotational speeds for various fibre mass concentrations.................................................................................... 132  xix  Figure 5.20: Modified mixing index as a function of dimensionless distance downstream for softwood and hardwood pulp suspensions at Cm = 1.0% Up = 0.5 m/s, R = 12, N = 400 rpm..................................................... 133 Figure 5.21: Modified mixing index as a function of dimensionless distance downstream for softwood and hardwood pulp suspensions with Cm = 0.5% Up = 2.0 m/s, R = 2 and N = 400 rpm........................................ 134 Figure 6.1: Tomographic images for air-water flow at Usl = 0.5 m/s, Usg = 0.11 m/s, Dr = 0.208: (a) without resistor adaptors (b) with resistor adaptors Locations of planes P1 to P8 are shown in Figure 3.2....................... 146 Figure 6.2: Gas mixing index as a function of dimensionless distance downstream for dispersed bubbles in air-water flow at Usl = 5.0 m/s and Dr = 0.208 for various superficial gas velocities. Locations of planes P2 to P8 are shown in Figure 3.2............................................................................ 148 Figure 6.3: Tomographic images for air-water flow at Usl = 5.0 m/s and Dr = 0.208 for: (a) Usg = 0.11 m/s (b) Usg = 0.44 m/s. Colour scales indicate gas holdup. Locations of planes P1 to P8 are shown in Figure 3.2.......... 149 Figure 6.4: Gas mixing index as a function of dimensionless distance downstream for air-water flow at Usl = 2.0 m/s and Dr = 0.208 for various superficial gas velocities...................................................................................... 151 Figure 6.5: Gas mixing index as a function of dimensionless distance downstream for the stratified air-water flow at Usl = 0.5 m/s and Dr = 0.208 for various superficial gas velocities.................................................................... 152 Figure 6.6: Gas mixing index as a function of dimensionless distance downstream for air-water flow at Usg = 0.11 m/s and Dr = 0.208 for various superficial liquid velocities. Letters designate flow regimes: DB: dispersed bubble; B: bubble; EB elongated bubble; S: stratified..................................... 153  xx  Figure 6.7: Vertical gas holdup profiles as a function of dimensionless distance downstream for the dispersed bubble air-water flow at Usl = 5.0 m/s, Usg = 0.11 m/s and Dr = 0.208.................................................................. 155 Figure 6.8: Vertical gas holdup profiles as a function of superficial gas velocity in airwater flow for Usl = 5.0 m/s, Dr = 0.208 and x/D = 22.1 (P8).............. 155 Figure 6.9: Correlation coefficient, R(r), as a function of separation distance, r, for the dispersed bubble air-water flow for Usl = 5.0 m/s, Dr = 0.208, x/D = 22.1 and various superficial gas velocities......................................... 156 Figure 6.10: Scale of segregation, Ls, as a function of dimensionless distance downstream for dispersed bubble air-water flow at Usl = 5.0 m/s and Dr = 0.208 for various superficial gas velocities...................................... 157 Figure 6.11: Comparison of effect of long vs. short injection tubes on gas mixing index for air-water flow for Usg = 0.11 m/s, Dj = 0.208 and various superficial liquid velocities. Letters designate flow regimes as in Figure 6.6...................................................................................................... 158 Figure 6.12: Comparison of effect of injection tube diameter on gas mixing index for air-water flow at Usg = 0.11 m/s for various superficial liquid velocities. Letters designate flow regimes as in Figure 6.6................................. 159 Figure 6.13: Average gas holdup as a function superficial gas and liquid velocities in air-water flow.................................................................................. 161 Figure 6.14: Average gas holdup for bubble and dispersed bubble flow patterns at Usl = 5.0 m/s for various superficial gas velocities and fibre mass concentrations and comparison with the volumetric flow ratio...........162 Figure 6.15: Gas mixing index as a function of dimensionless distance downstream for Usl = 5.0 m/s, Usg = 0.11 m/s and various fibre mass concentrations................................................................................... .164 xxi  Figure 6.16: Tomographic images for air-pulp suspension flow at Usl = 5.0 m/s and Usg = 0.11 m/s for: (a) Cm = 0.5% (b) Cm = 1.0% (c) Cm = 2.0% (d) Cm = 3.0%. Locations of planes P2, P4, P6 and P8 are shown in Figure 3.2. Tomographic images for air-water flow for the same flow conditions are illustrated in Figure 6.3a..................................................................... 165 Figure 6.17: Vertical gas holdup profiles as a function of fibre mass concentration for Usl = 5.0 m/s, Usg = 0.11 m/s at: (a) x/D = 12.2 and (b) x/D = 22.1.................................................................................................... 167 Figure 6.18: Scale of segregation as a function of dimensionless distance downstream for Usl = 5.0 m/s, Usg = 0.11 m/s and various fibre mass concentrations.................................................................................... 168 Figure 6.19: Gas mixing index as a function of dimensionless distance downstream for Usl = 2.0 m/s, Usg = 0.11 m/s and various fibre mass concentrations. Letters designate flow regimes as indicated in Figure 6.6................. 169 Figure 6.20: Tomographic images for air-pulp suspension flow at Usl = 2.0 m/s, Usg = 0.11 m/s for: (a) Cm = 0% (b) Cm = 1.0% (c) Cm = 2.0% (d) Cm = 3.0%. Locations of planes P2 – P5 are shown in Figure 3.2. Colour scales indicate gas holdups.......................................................................... 170 Figure 6.21: Scale of segregation, Ls, as a function of dimensionless distance downstream for Usl = 2.0 m/s, Usg = 0.11 m/s and various fibre mass concentrations.................................................................................... 171 Figure 6.22: Gas mixing index as a function of dimensionless distance downstream for Usl = 0.5 m/s, Usg = 0.11 m/s and various fibre mass concentrations. Letters designate flow regimes as indicated in Figure 6.6................. 172 Figure 6.23. Gas mixing index as a function of superficial liquid velocity for Usg = 0.11 m/s, x/D = 22.1 (plane P8) and various fibre mass concentrations.................................................................................... 174 xxii  Figure 6.24: Gas mixing index as a function of dimensionless distance downstream for Usg = 0.11 m/s, Dr = 0.208 and various superficial liquid velocities at: (a) Cm = 1.0% and (b) Cm = 3.0%. Letters designate flow regimes as indicated in Figure 6.6........................................................................ 176 Figure 6.25: Gas mixing index as a function of superficial gas velocity for Usl = 5.0 m/s, x/D = 22.1 (plane P8) and various fibre mass concentrations.... 178 Figure 6.26: Gas mixing index as a function of superficial gas velocity for Usl = 2.0 m/s, x/D = 22.1 (plane P8) and various fibre mass concentrations. EB indicates the elongated bubble flow regime....................................... 178 Figure 6.27: Gas mixing index as a function of superficial gas velocity for Usg = 0.5 m/s, x/D = 22.1 (plane P8) and various fibre mass concentrations. Letters identify flow regimes as indicated in Figure 6.6..................... 179 Figure 7.1: Gas mixing index and flow regimes as a function of dimensionless distance downstream, x/D, in air-water flow for N = 400 rpm and various superficial gas velocities for: (a) Usl = 3.0 m/s; (b) Usl = 1.0 m/s. P followed by a number designates sensor plane shown in Figure 3.5 Letters identify flow regimes as in Figure 6.6..................................... 186 Figure 7.2: Gas mixing index and flow regimes as a function of dimensionless distance downstream in air-water flow for Usg = 0.11 m/s, N = 400 rpm and various superficial liquid velocities. Letters identify flow regimes as in Figure 6.6....................................................................................... 187 Figure 7.3: Gas mixing index and flow regimes as a function of dimensionless distance downstream in air-water flow for Usg = 0.11 m/s and various impeller speeds at: (a) Usl = 3.0 m/s; (b) Usl = 0.5 m/s. Letters identify flow regimes as in Figure 6.6. // and  indicate parallel and normal to  flow..................................................................................................... 189  xxiii  Figure 7.4: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Usl = 3.0 m/s, Usg = 0.11 m/s, N = 600 rpm and various fibre mass concentrations. P2 – P8 identify planes, whereas other letters identify flow regimes....................................................... 190 Figure 7.5: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Usl = 0.5 m/s, Usg = 0.11 m/s, N = 600 rpm and various fibre mass concentrations...................................................... 192 Figure 7.6: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Usl = 1.0 m/s, Usg = 0.11 m/s, N = 600 rpm and various fibre mass concentrations...................................................... 193 Figure 7.7: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Usl = 0.5 m/s, virtually identical impeller speeds and various superficial gas velocities at: (a) Cm = 1.0%; (b) Cm = 2.0%................................................................................................... 195 Figure 7.8: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Usl = 1.0 m/s and Cm = 1.0%, with and without an impeller at: (a) Usg = 0.055 m/s; (b) Usg = 0.11 m/s..................................................................................................... 196 Figure 7.9: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Usl = 3.0 m/s and Cm = 1.0%, with and without an impeller at: (a) Usg = 0.11 and 0.33 m/s; (b) Usg = 0.44 m/s..................................................................................................... 198 Figure 7.10: Tomographic images for air-suspension flow for Usl = 3.0 m/s and Cm = 1.0% at: (a) Usg = 0.33 m/s; (b) Usg = 0.44 m/s. Locations of planes P1 – P8 are shown in Figure 3.5. Colour scales indicate gas holdup........ 200  xxiv  Figure 7.11: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 3.0%, Usl = 3.0 m/s and various superficial gas velocities: (a) with impeller; (b) without impeller......... 201 Figure 7.12: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Usg = 0.11 m/s, virtually identical impeller speeds and various superficial liquid/pulp velocities: (a) Cm = 1.0%; (b) Cm = 3.0%.......................................................................................... 203 Figure 7.13: Gas mixing index and flow regimes as a function of dimensionless distance downstream in air-suspension flow for Usl = 3.0 m/s, Usg = 0.11 m/s, Cm = 3.0% and various impeller speeds..................................... 205 Figure 7.14: Tomographic images for air-suspension flow for Usl = 3.0 m/s, Usg = 0.11 m/s and Cm = 3.0%: (a) tee mixer alone; (b) impeller rotating at 423 rpm. Locations of planes P2 – P8 are shown in Figure 3.5. Colour scales indicate gas holdup. ............................................................... 206 Figure 7.15: Gas mixing index and flow regimes as a function of dimensionless distance downstream in air-suspension flow for Usl = 1.0 m/s, Usg = 0.11 m/s and various impeller speeds at: (a) Cm = 0.5%; (b) Cm = 2.0%... 208 Figure 7.16: Gas mixing index and flow regimes as a function of dimensionless distance downstream in air-suspension flow for Usl = 0.5 m/s, Usg = 0.11 m/s, Cm = 3.0% and various impeller speeds..................................... 209 Figure 8.1: Schematic of D0 stage at Howe Sound Pulp and Paper..................... 214 Figure 8.2: ERT sensor plane (a) located at outlet of static mixer where pulp suspensions flow from right to left; (b) inserted between flanges downstream of static mixer................................................................ 216 Figure 8.3: ERT sensor design details (a) schematic of ERT sensor; (b) image of electrode............................................................................................ 217  xxv  Figure 8.4: Photographs of (a) old static mixer; (b) its mixer elements................ 220 Figure 8.5: Kappa factor representing chemical use before and after mixer installation.......................................................................................... 221 Figure 8.6: Temporal variation of electrical conductivity of pulp ahead of mixer and ClO2 solution....................................................................................... 223 Figure 8.7: Mixing index as a function of time after ClO2 valve was opened from shut-off position.................................................................................. 225 Figure 8.8: Tomographic images before and after ClO2 was introduced.............. 226 Figure 8.9: Temporal variation of mixing index as suspension flow rate changed from 390 to 225 L/s at Cm = 3.5%...................................................... 227 Figure 8.10: Tomographic images in time series when the suspension flow rate decreased from 390 to 264 L/s at Cm = 3.5%..................................... 228 Figure 8.11: Temporal variation of mixing index as consistency changed from 4.7% to 3.2% at a constant suspension flow rate of 325 L/s (Up = 1.1 m/s).................................................................................................... 229 Figure 8.12: Tomographic Images over time as the consistency decreased from 4.7% to 3.4% at a constant suspension flow rate of 325 L/s.............. 230 Figure A.1: Modified mixing index as a function of dimensionless distance downstream, x/D, with various jet-to-pipe velocity ratios and mainstream temperatures for water at Dr = 0.05……………………………………. 262 Figure A.2: Modified mixing index as a function of dimensionless distance downstream of injection for softwood pulp suspension at Cm = 0.5% and for water (w), Dr = 0.05 with various mainstream velocities for: (a) R ≈ 4; (b) R ≈ 6............................................................................................. 262  xxvi  Figure A.3: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Cm = 1.0% and for water (w), Dr = 0.05 with various mainstream velocities for: (a) R ≈ 3; (b) R ≈ 4......................................................................................................... 263 Figure A.4: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Cm = 1.0% and for water, Dr = 0.05 and various jet velocities for: (a) Up = 2.0 m/s in wall-source and jet-mixing modes; (b) Up = 1.0 m/s in jet-mixing and jet-mixing modes; (c) Up = 0.5 m/s in jet-mixing and jet-mixing modes.............. 263 Figure A.5: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Cm = 2.0% and for water (w), Dr = 0.05 with various mainstream velocities for: (a) R ≈ 4; (b) R ≈ 6......................................................................................................... 264 Figure A.6: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Cm = 3.0% and for water, Dr = 0.05 with various mainstream velocities for: (a) R ≈ 2; (b) R ≈ 3; (c) R ≈ 4................................................................................................... 264 Figure A.7: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Cm = 3.0% and for water, Dr = 0.05 and various jet velocities for: (a) Up = 2.0 m/s in wall-source and jet-mixing modes; (b) Up = 0.5 m/s in jet-mixing and jet-impaction modes................................................................................................ 265 Figure A.8: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension at Cm = 0.5% and for water, Dr = 0.05 with various mainstream velocities and similar jet-to-pipe velocity ratios of ~ 4........................................................................... 265 Figure A.9: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension at Cm = 0.5% and for water, xxvii  Dr = 0.05 and various jet velocities in jet-mixing and jet-impaction modes for: (a) Up = 1.0 m/s; (b) Up = 1.0 m/s..................................... 266 Figure A.10: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension at Cm = 1.0% and for water, Dr = 0.05 with various mainstream velocities for: (a) R ≈ 4; (b) R ≈ 6......................................................................................................... 266 Figure A.11: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension at Cm = 1.0% and for water, Dr = 0.05 and various jet velocities for: (a) Up = 3.0 m/s in wall-source and jet-mixing modes; (b) Up = 2.0 m/s in wall-source and jet-mixing modes; (c) Up = 1.0 m/s in jet-mixing and jet-impaction modes......... 267 Figure A.12: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension at Cm = 2.0% and for water, Dr = 0.05 with various mainstream velocities for: (a) R ≈ 2; (b) R ≈ 3; (c) R ≈ 4; (d) R ≈ 6................................................................................... 268 Figure A.13: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension at Cm = 2.0% and for water, Dr = 0.05 and various jet velocities in wall-source and jet-mixing modes for: (a) Up = 3.0 m/s; (b) Up = 2.0 m/s................................................. 268 Figure A.14: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension at Cm = 2.0% and for water, Dr = 0.05 and various jet velocities in jet-mixing and jet-impaction modes for (a) Up = 1.0 m/s; (b) Up = 0.5 m/s...................................... 269 Figure A.15: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension at Cm = 3.0% and for water, Dr = 0.05 with various mainstream velocities for: (a) R ≈ 2; (b) R ≈ 4......................................................................................................... 269  xxviii  Figure A.16: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension at Cm = 3.0% and for water, Dr = 0.05 and various jet velocities for: (a) Up = 3.0 m/s in wall-source and jet-mixing modes; (b) Up = 2.0 m/s in wall-source and jet-mixing modes; (c) Up = 1.0 m/s in jet-mixing and jet-impaction modes......... 270 Figure B.1: Modified mixing index as a function of dimensionless distance downstream of injection for water for Dr = 0.05 and various rotation speeds with: (a) Up = 1.0 m/s and R ≈ 6; (b) Up = 2.0 m/s and R ≈ 2......................................................................................................... 271 Figure B.2: Modified mixing index as a function of dimensionless distance downstream of injection for softwood pulp suspension with Cm = 0.5% and for water (w), Dr = 0.05, N = 400 rpm with various mainstream velocities and almost identical jet-to-pipe velocity ratios of ~6........... 271 Figure B.3: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension with Cm = 0.5% and for water (w), Up = 2.0 m/s, Dr = 0.05, N = 400 rpm and various jet-to-pipe velocity ratios..................................................................................... 272 Figure B.4: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension with Cm = 0.5% and for water (w), Dr = 0.05, Up = 1.0 m/s with various impeller speeds for: (a) R ≈ 4; (b) R ≈ 6............................................................................................. 272 Figure B.5: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension with Cm = 1.0% and for water (w), Dr = 0.05, N = 400 rpm with various mainstream velocities for: (a) R ≈ 4; (b) R ≈ 6...................................................................................... 272 Figure B.6: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension with Cm = 1.0% and for water  xxix  (w), Dr = 0.05, N = 400 rpm and various jet-to-pipe velocity ratios at: (a) Up = 2.0 m/s; (b) Up = 1.0 m/s; (c) Up = 0.5 m/s................................. 273 Figure B.7: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension with Cm = 1.0% and for water (w), Dr = 0.05, Up = 1.0 m/s with various impeller speeds for: (a) R ≈ 4; (b) R ≈ 6............................................................................................. 273 Figure B.8: Modified mixing index as a function of dimensionless distance downstream for softwood pulp with Cm = 2.0% and for water (w), Dr = 0.05, N ≈ 400 rpm and various velocity ratios for: (a) Up = 2.0 m/s; (b) Up = 1.0 m/s....................................................................................... 274 Figure B.9: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension with Cm = 2.0%, Dr = 0.05 and various impeller speeds for (a) Up = 1.0 m/s and R ≈ 4; (b) Up = 0.5 m/s and R ≈ 24................................................................................... 274 Figure B.10: Modified mixing index as a function of dimensionless distance downstream for softwood pulp with Cm = 3.0% and for water (w), Dr = 0.05, almost identical impeller speeds and various mainstream velocities at: (a) R ≈ 4; (b) R ≈ 8......................................................... 274 Figure B.11: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension with Cm = 3.0% and for water (w), Dr = 0.05, almost constant impeller rotation speed and various jet velocities at: (a) Up = 2.0 m/s; (b) Up = 1.0 m/s; (c) Up = 0.5 m/s....... 275 Figure B.12: Modified mixing index as a function of dimensionless distance downstream for softwood pulp with Cm = 3.0% and for water (w), Dr = 0.05 and various impeller speeds for: (a) Up = 1.0 m/s and R ≈ 4; (b) Up = 0.5 m/s and R ≈ 8............................................................................ 275  xxx  Figure B.13: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension with Cm = 0.5% and for water (w), Dr = 0.05, N ≈ 400 rpm and R ≈ 4, with various mainstream velocities............................................................................................ 276 Figure B.14: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension with Cm = 0.5% and for water (w), Dr = 0.05, N ≈ 400 rpm and various velocity ratios at: (a) Up = 2.0 m/s; (b) Up = 0.5 m/s.................................................................... 276 Figure B.15: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension with Cm = 0.5% and for water (w), Dr = 0.05, Up = 1.0 m/s and R ≈ 6 with various impeller speeds................................................................................................ 276 Figure B.16: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension with Cm = 1.0% and for water (w), Dr = 0.05, N = 400 rpm and various mainstream velocities at: (a) R ≈ 4; (b) R ≈ 6............................................................................. 277 Figure B.17: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension with Cm = 1.0% and for water (w), Dr = 0.05, N ≈ 400 rpm and various velocity ratios at: (a) Up = 2.0 m/s; (b) Up = 1.0 m/s; (c) Up = 0.5 m/s......................................... 277 Figure B.18: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension with Cm = 1.0% and for water (w), Dr = 0.05, Up = 1.0 m/s and various impeller speeds at: (a) R ≈ 4; (b) R ≈ 6; (c) R ≈ 8....................................................................... 278 Figure B.19: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension with Cm = 2.0% and for water (w), Dr = 0.05, almost identical impeller rotation speeds and various mainstream velocities at: (a) R ≈ 4; (b) R ≈ 6........................ 278 xxxi  Figure B.20: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension with Cm = 2.0% and for water (w), Dr = 0.05, almost identical impeller rotation speeds and various velocity ratios at: (a) Up = 2.0 m/s; (b) Up = 1.0 m/s; (c) Up = 0.5 m/s..................................................................................................... 279 Figure B.21: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension with Cm = 2.0% and for water (w), Dr = 0.05 and various impeller speeds at: (a) Up = 2.0 m/s and R ≈ 2; (b) Up = 1.0 m/s R ≈ 4....................................................... 279 Figure B.22: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension with Cm = 2.0% and for water (w), Dr = 0.05 and various impeller speeds at: (a) Up = 1.0 m/s and R ≈ 6; (b) Up = 0.5 m/s R ≈ 8....................................................... 280 Figure B.23: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension with Cm = 3.0% and for water (w), Dr = 0.05, almost identical impeller rotation speeds and various mainstream velocities at: (a) R ≈ 4; (b) R ≈ 6........................ 280 Figure B.24: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension with Cm = 3.0% and for water (w), Dr = 0.05, almost identical impeller rotation speeds and various jet velocities at: (a) Up = 2.0 m/s; (b) Up = 1.0 m/s; (c) Up = 0.5 m/s..................................................................................................... 281 Figure B.25: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension with Cm = 3.0% and for water (w), Dr = 0.05, Up = 1.0 m/s and R ≈ 6 with various impeller speeds................................................................................................ 281 Figure C.1: Gas mixing index and flow regimes as a function of dimensionless distance downstream for air-water flow for various superficial gas xxxii  velocities at: (a) Usl = 1.0 m/s; (b) Usl = 3.0 m/s; (c) Usl = 4.0 m/s. Letters identify flow regimes as in Figure 6.6................................................. 282 Figure C.2: Gas mixing index and flow regimes as a function of dimensionless distance downstream for air-water flow for various superficial liquid velocities at: (a) Usg = 0.22 m/s; (b) Usg = 0.33 m/s; (c) Usg = 0.44 m/s. Letters identify flow regimes as in Figure 6.6..................................... 283 Figure C.3: Gas mixing index and flow regimes as a function of dimensionless distance downstream for air-water flow at Usg = 0.11 m/s for various superficial liquid velocities and mainstream temperatures................. 283 Figure C.4: Vertical gas holdup profiles as a function of dimensionless distance downstream for air-water flow at Usg = 0.11 m/s for: (a) Usl = 2.0 m/s; (b) Usl = 3.0 m/s; (c) Usl = 4.0 m/s............................................................ 284 Figure C.5: Gas mixing index as a function of dimensionless distance downstream for Cm = 0.5% and various superficial gas velocities at: (a) Usl = 0.5 m/s; (b) Usl = 1.0 m/s. Letters identify flow regimes as in Figure 6.6......... 284 Figure C.6: Gas mixing index as a function of dimensionless distance downstream for Cm = 0.5% and various superficial gas velocities at: (a) Usl = 2.0 m/s; (b) Usl = 3.0 m/s; (c) Usl = 4.0 m/s; (d) Usl = 5.0 m/s. Letters identify flow regimes as in Figure 6.6..................................................................... 285 Figure C.7: Gas mixing index as a function of dimensionless distance downstream for Cm = 0.5% and various superficial liquid velocities at: (a) Usg = 0.11 m/s; (b) Usg = 0.22 m/s; (c) Usg = 0.33 m/s; (d) Usg = 0.44 m/s. Letters identify flow regimes as in Figure 6.6................................................. 286 Figure C.8: Gas mixing index as a function of dimensionless distance downstream for Cm = 1.0% and various superficial gas velocities at: (a) Usl = 0.5 m/s; (b) Usl = 1.0 m/s; (c) Usl = 2.0 m/s; (d) Usl = 3.0 m/s; (e) Usl = 4.0 m/s; (f) Usl = 5.0 m/s. Letters identify flow regimes as in Figure 6.6............... 287 xxxiii  Figure C.9: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 2.0% and various superficial gas velocities at: (a) Usl = 0.5 m/s; (b) Usl = 1.0 m/s; (c) Usl = 2.0 m/s; (d) Usl = 3.0 m/s; (e) Usl = 4.0 m/s; (f) Usl = 5.0 m/s. Letters identify flow regimes as in Figure 6.6..................................................................... 288 Figure C.10: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 3.0% and various superficial gas velocities at: (a) Usl = 0.5 m/s; (b) Usl = 1.0 m/s; (c) Usl = 2.0 m/s; (d) Usl = 4.0 m/s; (e) Usl = 5.0 m/s. Letters identify flow regimes as in Figure 6.6...................................................................................................... 289 Figure D.1: Gas mixing index and flow regimes as a function of dimensionless distance downstream in air-water flow for N = 400 rpm and various superficial gas velocities at: (a) Usl = 0.5 m/s; (b) Usl = 2.0 m/s. Letters identify flow regimes as in Figure 6.6................................................. 290 Figure D.2: Gas mixing index and flow regimes as a function of dimensionless distance downstream in air-water flow for Usl = 2.0 m/s, Usg = 0.11 m/s and various impeller speeds. Letters identify flow regimes as in Figure 6.6...................................................................................................... 290 Figure D.3: Gas mixing index and flow regimes as a function of dimensionless distance downstream in air-water flow for N = 400 rpm and various superficial liquid velocities at (a) Usg = 0.22 m/s; (a) Usg = 0.33 m/s; (a) Usg = 0.44 m/s. Letters identify flow regimes as in Figure 6.6............ 291 Figure D.4: Average gas holdup as a function superficial gas velocity for various superficial liquid velocities and mainstream temperatures in air-water flow..................................................................................................... 291 Figure D.5: Gas mixing index as a function of dimensionless distance downstream in air-water flow for Usl = 2.0 m/s, Usg = 0.44 m/s, N = 400 rpm and  xxxiv  various mainstream temperatures. Letters identify flow regimes as in Figure 6.6........................................................................................... 292 Figure D.6: Gas mixing index and flow regimes as a function of dimensionless distance downstream in air-water flow for Usg = 0.11 m/s, N = 400 rpm and various superficial liquid velocities and mainstream temperatures...................................................................................... 292 Figure D.7: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 0.5%, N = 400 rpm and various superficial gas velocities at: (a) Usl = 0.5 m/s; (b) Usl = 1.0 m/s......... 292 Figure D.8: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 0.5%, N = 400 rpm and various superficial gas velocities at: (a) Usl = 2.0 m/s; (b) Usl = 3.0 m/s......... 293 Figure D.9: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 0.5%, N = 400 rpm and various superficial liquid/pulp velocities at: (a) Usg = 0.11 m/s; (b) Usg = 0.22 m/s; (c) Usg = 0.33 m/s; (d) Usg = 0.44 m/s. Letters identify flow regimes as in Figure 6.6.................................................................................. 293 Figure D.10: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 0.5%, Usg = 0.11 m/s and various impeller rotation speeds at: (a) Usl = 0.5 m/s; (b) Usl = 2.0 m/s; (c) Usl = 3.0 m/s. Letters identify flow regimes as in Figure 6.6....................... 294 Figure D.11: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 1.0%, virtually identical impeller speeds and various superficial gas velocities at: (a) Usl = 1.0 m/s; (b) Usl = 2.0 m/s..................................................................................................... 294 Figure D.12: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 1.0%, virtually identical impeller speeds xxxv  and various superficial liquid/pulp velocities at (a) Usg = 0.22 m/s; (b) Usg = 0.33 m/s; (c) Usg = 0.44 m/s. Letters identify flow regimes as in Figure 6.6........................................................................................... 295 Figure D.13: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 1.0%, Usg = 0.11 m/s and various impeller speeds at: (a) Usl = 0.5 m/s; (b) Usl = 1.0 m/s. Letters identify flow regimes as in Figure 6.6............................................................. 295 Figure D.14: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 1.0%, Usg = 0.11 m/s and various impeller speeds at: (a) Usl = 2.0 m/s; (b) Usl = 3.0 m/s. Letters identify flow regimes as in Figure 6.6............................................................. 296 Figure D.15: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 2.0%, virtually identical impeller speeds and various superficial liquid/pulp velocities at: (a) Usg = 0.11 m/s; (b) Usg = 0.22 m/s; (c) Usg = 0.33 m/s,; (d) Usg = 0.44 m/s....................... 296 Figure D.16: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 2.0%, almost identical impeller speeds and various superficial gas velocities at: (a) Usl = 1.0 m/s; (b) Usl = 2.0 m/s. Letters identify flow regimes as in Figure 6.6............................. 297 Figure D.17: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 2.0%, Usg = 0.11 m/s and various impeller speeds at (a) Usl = 0.5 m/s; (b) Usl = 2.0 m/s; (c) Usl = 3.0 m/s..................................................................................................... 297 Figure D.18: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 3.0%, almost identical impeller speeds and various superficial gas velocities at: (a) Usl = 0.5 m/s; (b) Usl = 1.0 m/s; (c) Usl = 2.0 m/s. Letters identify flow regimes as in Figure 6.6...................................................................................................... 298 xxxvi  Figure D.19: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 3.0%, Usg = 0.11 m/s and various impeller speeds at: (a) Usl = 1.0 m/s; (b) Usl = 2.0 m/s. Letters identify flow regimes as in Figure 6.6............................................................ .298 Figure D.20: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Cm = 3.0%, virtually identical impeller speeds and various superficial liquid/pulp velocities at: (a) Usg = 0.22 m/s; (b) Usg = 0.33 m/s; (c) Usg = 0.44 m/s. Letters identify flow regimes as in Figure 6.6........................................................................................... 299  xxxvii  Nomenclature Ci  local salt concentration in mixture, kg/m3  Cj  salt concentration in side stream or jet flow, kg/m3  Cm  suspension mass concentration or consistency, %  Cp  salt concentration in main stream or pipe flow, kg/m3  Cv  suspension volumetric concentration, dimensionless  C  average salt concentration of mixture, kg/m3  d  fibre diameter, m  D  pipe diameter, m  Dj  injection tube diameter, m  Dr  jet-to-pipe diameter ratio, dimensionless  Is  intensity of segregation, dimensionless  L  fibre length, m  Lj  injection tube length, m  Ls  scale of segregation representing relative size of gaseous entities, m  M  mixing index, %  MFS  mixing index for fully-segregated flow, %  Mg  gas mixing index, %  Mm  measured mixing index, %  Ms  system mixing index measured in absence of tracer, %  M’  modified mixing index, %  n  number of image pixels, dimensionless  xxxviii  ne  number of electrodes  N  impeller rotation speed, rpm  Nc  crowding number, dimensionless  Nm  number of unique measurements  Qj  volumetric flow rate of side stream, m3/s  Qp  volumetric flow rate of main stream, m3/s  r  pixel distance between two points in mixture, m  R  jet-to-pipe mean velocity ratio, i.e. Uj/Up, dimensionless  r0  value of r at which R(r) reaches zero, m  R(r)  coefficient of correlation, dimensionless  Rej  jet Reynolds number, dimensionless  Rep  pipe Reynolds number, dimensionless  t  time, s  Up  mainstream or pipe velocity, m/s  Uj  side-stream or jet velocity, m/s  Usg  superficial gas velocity, m/s  Usl  superficial liquid/pulp velocity, m/s  x  distance downstream of injection, m  y1  electrical conductivity of continuous phase, mS/cm  y2  electrical conductivity of dispersed phase, mS/cm  ymc,i  local mixture electrical conductivity from ERT measurement, mS/cm  y  average electrical conductivity, mS/cm  xxxix  Greek letters    overall mixing index, m-1   H/L  friction loss, m water/100 m pipe  g,i  local gas holdup from ERT measurement, dimensionless  g  average gas holdup in a cross-sectional plane, dimensionless    standard deviation of conductivity in tomographic image, mS/cm   FS  standard deviation for fully segregated flow, mS/cm  g  modified standard deviation, dimensionless  g, FS  standard deviation of local gas holdup for fully-segregated flow, dimensionless  m  standard deviation of local gas holdup, dimensionless  s  system standard deviation of local gas holdup measured in absence of gas, dimensionless   2 (0)  conductivity variance before tracer addition, (mS/cm)2    fibre coarseness, kg/m  Abbreviations B  bubble flow  DB  dispersed bubble flow  EB  elongated bubble flow  S  stratified flow xl  SL  slug flow  W  wavy flow  xli  Acknowledgements I would like to express my sincere gratitude and appreciation to my supervisors Prof. John R. Grace and Prof. C. Jim Lim for their invaluable guidance and support throughout my thesis. I am grateful for their supervision with patience, insightful scientific knowledge and motivation. I would also like to sincerely acknowledge my late supervisor Prof. Chad P.J. Bennington for introducing me to pulp and paper science and technology, for initiating this research project and for giving me guidance, knowledge and encouragement. Without Prof. Bennington, this work would not have been possible. I am grateful to Prof. Richard J. Kerekes for technical assistance, and to Profs. James A. Olson and D. Mark Martinez for providing access to the flow loop facility and liquid injection system. I also appreciate Profs. Richard J. Kerekes and David G. Dixon for their willingness to serve on the examining committee. I am grateful for the financial support from the University of British Columbia and NSERC. I greatly appreciate Joanne Dean for her excellent support of the NSERC grant application. I acknowledge Gerry Pageau and colleagues at Howe Sound Pulp and Paper Ltd. for fabricating and installing the sensor array and for their support during mill trials. I am also grateful to Gerry Pageau for providing mill data for the design of experiments and equipment and for data analysis, and for his helpful advice and knowledge in pulp and paper science.  xlii  I thank Drs. Mannish Bhole and Subhashini Vashisth for their technical advice. I am grateful to George Soong and Tim Paterson for their assistance in the laboratory. I also thank all members in the Dynamic Mixing Group and colleagues from the Pulp and Paper Centre for initiating many ideas and for their knowledge and encouragement. Thanks are due to Doug Yuen, Serge Milaire and members of the CHBE workshop and stores for technical assistance. I also thank members in the CHBE administration, in particular Helsa Leong, for their support. From the bottom of my heart, I wish to express my appreciation to my parents, my brothers and sisters, and my aunts for their continuous love, inspiration and support.  xliii  1. Introduction1 Mixing is an essential unit operation in the pulp and paper industry. It is used for blending pulp in stock chests, controlling consistency in pulp processing, attenuating variations in consistency before paper machines, and mixing chemicals into pulp suspensions for bleaching. Without good mixing, operating costs can increase due to unstable production and poor product quality. For example, poor mixing in a bleach plant leads to poor chemical contacting, resulting in poor product quality such as lower pulp brightness and cleanliness. This is often offset by adding extra chemicals, increasing production costs and adversely affecting product strength. The benefits of improved mixing have been known for many years. Mill experience and laboratory studies have shown that improved mixing leads to reduced chemical usage and more uniform product quality in pulp bleaching (e.g. Elliott and Farr, 1973; Torregrossa, 1983; Backlund et al., 1987). Several pulp mills obtained returns on investment in as little as 3 months after improving mixing efficiency (Bennington, 2004). Although several theoretical and experimental studies have been reported in the literature on mixing of pulp suspensions in stirred tanks, there are few studies on in-line mixing of these suspensions. In bleach plants, pre-distribution of chemicals  1  Part of this chapter was published as Yenjaichon, W., Pageau G., Bhole, M., Bennington, C.P.J,  Grace, J.R, 2011. Assessment of mixing quality for an industrial pulp mixer using electrical resistance tomography. Can. J. Chem. Eng. 89 (5), 996–1004.  1  in pulp suspensions before entering tower reactors is essential. In-line mixers are often used to ensure efficient contact between chemicals and pulp to achieve effective lignin removal and optimal use of the bleaching agent. Chemical injection in pulp mixing is commonly used as a pre-distributor to disperse chemicals into pulp suspensions ahead of, or inside, various mixers, involving static mixers, peg mixers and high-shear mixers. In modern bleaching processes, in-line mixers have replaced continuous stirred vessels due to increased energy dissipation rate per unit volume (Bennington, 1996). For low-consistency (fibre mass concentration) operations, such as the first chlorine dioxide bleaching (D0) stage (liquid-suspension mixing), in-line static mixers have been widely used. High-shear mixers have often been utilized in medium-consistency applications such as chlorine dioxide bleaching (liquid-suspension), oxygen delignification (gas-suspension) and oxidative alkaline extraction (gas-suspension) stages. In-line mixers are also used for steam mixing (gas-suspension) ahead of chemical mixers. Understanding in-line mixing of pulp suspensions can lead to criteria and guidance for the design of in-line mixers and for process optimization. This study focuses on in-line mixing of both liquid- and gas-suspensions. At present, there is still no effective method for determining mixing quality on the industrial scale. Most mixing assessment techniques are intrusive, tedious, and time consuming. Temperature profiling the surface of a process pipe does not interfere with the process, but only measures mixing quality at the periphery of the process, and a suitable temperature difference between the pulp suspension and  2  the added chemical must exist. Therefore, a more suitable way to assess mixing quality is desirable. Electrical resistance tomography (ERT) is a non-intrusive technique utilized extensively to quantify mixing in various processes (e.g. Williams et al., 1996; Stephenson et al., 2007; Hosseini et al., 2010). ERT measures the distribution of conductivity in the region of interest from voltage measurements at the vessel periphery (Mann et al., 1997). In the laboratory, ERT can be used to determine the progress of mixing over a range of conditions. In industry, if a suitable conductivity difference between the suspension and added chemical exists, the spatial homogeneity of conductivity can be measured, and mixing quality can be quantified.  1.1 Thesis Objectives In this study, ERT is applied to assess mixing quality in pulp suspensions on both laboratory and industrial scales. On the laboratory scale, liquid distribution and gas dispersion in pulp suspension flow are investigated downstream of in-line mixers, including jet mixers and a mechanical mixer, based on novel mixing indices, derived from the standard deviation of cross-sectional image pixel values. On the industrial scale, the same technique is implemented to evaluate mixing quality for an industrial pulp mixer over a range of operating conditions. The main objectives of this study are: 1. To develop criteria for optimizing in-line liquid jet mixing in non-Newtonian pulp suspensions and provide better understanding of jet mixing behaviour 3  for different flow regimes, jet penetration and fibre properties, based on a novel liquid mixing index (addressed in Chapter 4). 2. To examine gas dispersion in horizontal pulp-fibre-suspension flow for different gas and suspension flow regimes, based on a novel gas mixing index (see Chapter 6). 3. To evaluate the quality of liquid/gas mixing into pulp suspension flow downstream of an in-line mechanical mixer over a range of operating conditions and establish optimum operating conditions (Chapters 5 and 7). 4. To test the implementation of ERT as a real-time mixing assessment tool in an industrial pulp bleaching stage to monitor temporal variations of process conditions and evaluate their effects on mixing quality for process optimization (addressed in Chapter 8).  1.2 Thesis Organization The present chapter discusses the motivation and main objectives for this study. Chapter 2 includes a brief introduction of pulp bleaching processes, mixing equipment and the benefits of improved mixing. General methods to measure mixing quality and past techniques used to assess mixing in pulp suspensions on both laboratory and industrial scales are also discussed. This chapter also covers principles of electrical resistance tomography and its applications to various processes, including pulp and paper operations. 4  Chapter 3 describes details of the experimental equipment including the pilotscale flow loop facility, test section, ERT system, liquid injection system and gas injection system. We also provide the relevant properties of softwood and hardwood kraft pulps used in this study, describe the test procedure, and introduce mixing indices for both liquid-suspension and gas-suspension mixing. Chapter 4 discusses significant differences in jet mixing of a Newtonian fluid (water) in the turbulent flow regime and non-Newtonian pulp suspensions. Effects of injection tube length, mainstream and jet velocities, and fibre mass concentration on mixing quality for water, softwood and hardwood pulp suspensions are evaluated. Influences of fibre-turbulence interactions in modifying turbulent structures in the bulk are also discussed. This chapter is based on a journal paper already published (Yenjaichon, W., Grace, J.R., Lim, C.J., Bennington, C.P.J., 2012. In-line jet mixing of liquid-pulp-fibre suspensions: Effect of concentration and velocities. Chem. Eng. Sci. 75, 167–176) and a paper accepted for publication (Yenjaichon, W., Grace, J.R., Lim, C.J., Bennington, C.P.J., 2012. In-line jet mixing of liquid-pulp-fiber suspensions: Effect of fiber properties, flow regime and jet penetration, AIChE J.). Chapter 5 describes mixing behaviour of liquid injection into water or pulp suspensions for an in-line mechanical mixer over a range of operating conditions including mainstream and jet velocities, fibre mass concentrations, rotational speeds and fibre types. Mixing is compared with that for the tee alone. Decaying turbulence accompanied by fibre reflocculation downstream of the impeller is also discussed. This chapter is based on a paper accepted for publication (Yenjaichon, W., Grace, J.R., Lim, C.J., Bennington, C.P.J., 2012. Pilot-scale examination of  5  mixing liquid into pulp fiber suspensions in the presence of an in-line mechanical mixer, Ind. Eng. Chem. Res.). Gas dispersion in horizontal pipe flow downstream of a 90-tee mixer for both water and softwood kraft pulp suspensions is investigated in Chapter 6. Both a novel gas mixing index and the scale of segregation are deployed to evaluate mixing in water and pulp suspension flow for various flow regimes. Effects of injection tube length and diameter, superficial gas and liquid/pulp velocities, and fibre mass concentration on mixing quality are also determined. This chapter is based on two papers accepted for publication (Yenjaichon, W., Grace, J.R., Lim, C.J., Bennington, C.P.J., 2012. Characterisation of gas mixing in water and pulpsuspension flow based on electrical resistance tomography, Chem. Eng. J.; and Yenjaichon, W., Grace, J.R., Lim, C.J., Bennington, C.P.J., 2012. Gas dispersion in horizontal pulp-fibre-suspension flow, Int. J. Multiphase Flow). Chapter 7 is concerned with dispersion of gas in pulp suspensions downstream of an in-line mechanical mixer. Influences of buoyancy, fibre network and impeller speed on gas mixing are studied. This chapter is very similar to a paper submitted for publication (Yenjaichon, W., Grace, J.R., Lim, C.J., Bennington, C.P.J., 2012. Gas dispersion in pulp-suspension flow in the presence of an in-line mechanical mixer). Chapter 8 presents the results of implementation of ERT to assess mixing performance of a static mixer in an industrial chlorine dioxide bleaching stage. Influences of the chemical flow rate, suspension flow rate and suspension mass concentration on mixing quality are delineated. The results are compared with data  6  in the literature based on other measurement techniques for similar mixer installations. This chapter is based on a published journal paper (Yenjaichon, W., Pageau G., Bhole, M., Bennington, C.P.J, Grace, J.R, 2011. Assessment of mixing quality for an industrial pulp mixer using electrical resistance tomography. Can. J. Chem. Eng. 89(5), 996–1004). Chapter 9 provides conclusions, contributions to knowledge, potential applications, limitations of the research and recommendations for future research.  7  2. Background and Literature Review 2.1 Pulp Bleaching Processes Pulp bleaching is a chemical process used to remove coloured components that contaminate pulp fibres. Lignin, a highly coloured component, is removed in a sequence of steps including mixing chemicals with the pulp suspension, allowing the chemicals to react with the lignin, and washing the suspension. In a typical bleaching stage, pulp suspension from a previous stage is washed to remove reacted and dissolved substances. The suspension then enters a steam mixer to increase the reaction temperature before passing through a chemical mixer where a bleaching chemical is added. The pulp suspension then enters a tower reactor with sufficient time to complete the reaction. Finally, it is washed before being transferred to the next process. The details of bleaching processes are provided by Reeve (1996). A process diagram of a bleaching stage is shown in Figure 2.1. Both liquid and gas serve as bleaching chemicals. Liquid chemicals include chlorine dioxide, hydrogen peroxide, hypochlorite and sodium hydroxide, whereas gaseous chemicals include oxygen and ozone. Typical operating conditions for bleaching stages at Howe Sound Pulp and Paper Ltd. are summarized in Tables 2.1 and 2.2. These conditions are used in this thesis to design laboratory experiments and to evaluate the application of electrical resistance tomography (ERT) to assess mixing quality on an industrial scale.  8  Figure 2.1: Process diagram of a typical bleaching stage (adapted from Bennington, 2004)  Table 2.1: Typical operating conditions of first chlorine dioxide (D0) and chlorine dioxide D1 bleaching stages (liquid-suspension applications) at Howe Sound Pulp and Paper (HSPP) Ltd. (Data courtesy of the company)  Typical operating conditions Production rate (ADt/d) Pulp consistency (%) ClO2 on pulp (%) ClO2 strength (g/L) ID of mainstream pipe (mm) ID of side-stream pipe (mm) Pressure at flange location (kPa) Mainstream mean velocity (m/s) Side-stream mean velocity (m/s) Electrical conductivity (mS/cm) - Pulp suspension - ClO2 solution a  D0 stage 800 – 1,400 3.4 – 3.8 1.0 – 2.5 10 – 10.5 610 89 360 0.8 – 1.4 2.7 – 4.7  D1 stage 400 – 800 10.5 – 11.5 0.85 – 1.2 10 – 10.5 254 43 600 0.7 – 1.4 2.7 – 5.4  4.8 1.76  NAa 1.76  not available. 9  Table 2.2: Typical operating conditions of oxygen delignification (O) and oxidative alkaline extraction (EO) stages (gas-suspension applications) at HSSP (Data courtesy of the company) Typical operating conditions Production rate (ADt/d) Pulp consistency (%) O2 on pulp (%) O2 feed pressure (kPa) Pressure after O2 mixer (kPa) Temperature after O2 mixer (oC) ID of mainstream pipe (mm) ID of side-stream pipe (mm) Mainstream mean velocity (m/s) Gas volume fraction  O stage 800 – 1,500 8.5 – 9.5 1.0 – 2.5 1,400 800 90 218 74 2.3 – 4.8 0.09 – 0.22  EO stage 400 – 800 10 –11 0.5 1,400 300 90 175 25 1.5 – 3.3 0.14 – 0.15  2.2 Mixing Equipment in Pulp Bleaching Processes Various types of mixers have been used for pulp bleaching. Mixer design has improved over time as a result of changes in the requirements of bleaching processes and increased knowledge of pulp rheology. For example, static mixers have replaced stirred tanks in low-consistency operations, due to greater power dissipation. Mixers are categorized according to fibre mass concentration or intended consistency of pulp suspension (Cm), including low, medium and high consistency, with details provided by Bennington (1996, 2004).  10  2.2.1 Low-consistency mixers Creating motion is relatively easy at low fibre mass concentrations (Cm< 5%) since the yield stress is relatively low and abundant free water exists in suspension. Mixers utilized at these concentrations include stirred/agitated vessels, static mixer and dynamic high-shear mixers. Agitated vessels were used in the past, and have been replaced by static mixers and dynamic high-shear mixers since increased energy dissipation improves bleaching and reduces chemical usage. In bleach plants, static mixers are commonly used for low-consistency applications such as the first chlorine dioxide (D0) stage. They are also utilized in a few medium-consistency applications, but are unable to generate enough intense shear to generate turbulence (Bennington, 1996). Static mixers or in-line motionless mixers, consisting of a series of fixed mixing elements along the pipe, produce mixing by dividing and recombining the flow, by generating turbulence in the flow, or by combining both mechanisms. The energy required for mixing comes from the pressure drop when fluids are forced through mixing elements by the pumping system. Static mixers are used in both laminar and turbulent processes. For viscous liquids in the laminar regime, mixing is achieved by dividing the flow into substreams, distributing the substreams radially, and recombining them in an ordered sequence (Etchells and Meyer, 2004). This reduces the volume of individual regions of non-uniformity by continual rearrangement of the mixer contents, a major aspect of static mixer performance. Thus, mixing can be improved by increasing the number of mixing elements along the pipe (Godfrey, 1985; Edwards, 1985). For low-viscosity fluids, turbulent flow is usually established, and 11  mixing is achieved by increasing the level of turbulence. Uniformity is then achieved more easily than for laminar mixing.  2.2.2 Medium-consistency mixers The range of medium consistency is from 8 to 16%. Medium consistency mixers include peg mixers, high-shear or high-intensity mixers, medium consistency pumps, and valve and pipe expansions. High-shear mixers are now common in medium-consistency operations such as chlorine dioxide bleaching and oxygen delignification. In bleaching processes, jet mixers are usually utilized in conjunction with peg mixers and high-shear mixers. Peg mixers are tubular vessels consisting one or two shafts with attached pegs. Chemicals are introduced to pulp suspensions through injection ports located upstream of, or inside, the mixer. Multiple injection ports are common for both chemical and steam mixing. For high-shear mixers, chemicals are injected into pulp suspensions during passage through zones of high shear. These zones are created by high rotational speeds across narrow gaps, leading to high power dissipation per unit volume. At these concentrations, suspensions reflocculate rapidly downstream of high-shear zones. A common design criterion is to disrupt fibre flocs during contacting with bleaching chemicals by providing sufficient energy to “fluidize” the suspension in the mixing zone.  12  2.2.3 High-consistency mixers At high consistencies (Cm> 20%), pulp suspensions have little or no free water. Mixing is achieved by creating surface area within the suspension to increase contact between fibres and bleaching chemicals, e.g. by disrupting the flocs using strong shearing forces. High-consistency mixers include Schredder type and Kneader type mixers (Bennington, 2004).  2.3 Benefits of Improved Mixing in Pulp Bleaching In typical pulp bleaching, chemicals are added to pulp suspensions to react with lignin, distributed in the fibres. In order to obtain high-quality bleached pulp, good mixing is required to distribute sufficient chemicals to all the fibres. Underbleaching leads to poor pulp quality e.g. low pulp brightness, whereas overbleaching increases production costs and decreases pulp viscosity and strength. Several studies have been carried out to evaluate the benefits of improved mixing (Atkinson and Partridge, 1966; Elliott and Farr, 1973; Torregrossa, 1983; Kolmodin, 1984; Pattyson, 1984; Sin, 1984; Backlund et al., 1985; Berry, 1985; Backlund and Parming, 1987; Robitaille, 1987). In 1990, a CPPA survey reported the chemical savings from replacing existing mixers with high-shear mixers in the chlorine and chlorine dioxide stages (Berry, 1990). Average savings were 7 – 8 kg Cl2/t for the chlorine stage and 2.0 – 2.4 kg ClO2/t for the chlorine dioxide stage.  13  Benefits of improved mixing can be predicted by laboratory data and computer simulations. The bleaching responses under ideal mixing conditions in the laboratory need to be known so that pulp properties can be predicted for various mixing conditions. Torregrossa (1983) evaluated benefits of improved mixing using laboratory simulation based on chlorine dioxide charge deviation, defined as percent deviation from the average value, and a rectangular (or continuous uniform) distribution of the added chemical. Results from the laboratory simulation agreed with mill results, showing that improved mixing enhanced pulp properties. Backlund and Parming (1987) determined the effects of non-homogeneity on fibre and shive bleaching in the chlorine dioxide bleaching (D1) stage for various mixing conditions based on the coefficient of variation. The simulation was based on kinetic, stoichiometric and buffering equations for fibre and shive bleaching, and various statistical distributions of the applied chemical. The results indicated that improved mixing provided better brightness and lower shive content.  2.4 Mixing Scales The objective of mixing is to produce a uniform distribution of added chemicals in pulp suspensions, eliminating non-uniformities over a range of length scales. Non-uniformity scales can be defined according to the distance of relative motion required to eliminate concentration variations (Bennington, 1996). The largest of these is the macroscale, taken as 10 mm or larger, but it could be as large as the dimensions of mixing vessel. For this scale, mixing is achieved primarily by 14  bulk movement of fluid over relatively large distances. The next scale is the fibrescale, representing non-uniformity having a size from fibres to flocs, 0.05 - 10 mm. The variations for this scale can be reduced by laminar and turbulent shear and by diffusion. The smallest scale is the microscale, representing non-uniformity approaching molecular level. Uniformity at this scale is attained by molecular diffusion, aided by small-scale fluid motion.  2.5 Measurement of Mixing Quality Several studies have been carried out to characterize mixing quality. One of the most comprehensive methods was proposed by Danckwerts (1952). The mixing of two mutually soluble liquids was visualized as one is added to the other. The added liquid is broken down into clumps and distributed uniformly throughout a mixture by mechanical mixing. The clumps are dispersed and reduced in size until they can no longer be reduced by mechanical mixing. At the same time, molecular diffusion occurs at the boundary of the clumps where there is a concentration difference between the clumps and surrounding fluid. This process continues even if mechanical mixing stops. Two parameters were proposed by Danckwerts (1952) to measure the mixing quality based on the statistical theory of turbulence. The first parameter is the scale of segregation (Ls), a measure of the size of segregation regions or “clumps”, obtained from the area under the curve of the coefficient of correlation plotted as a function of separation distance. The second is the intensity of 15  segregation (Is), providing a measure of concentration difference between clumps and surrounding fluid. Several subsequent studies have used these two parameters to quantify mixing. Paterson and Kerekes (1985) evaluated mixing quality in pulp fibre suspensions and established standards for good and poor mixing based on Is and Ls of the residual chlorine content. They also measured microscale mixing in six mill chlorination mixers in bleaching processes of five kraft mills (Paterson and Kerekes, 1986). Pulp samples were collected downstream of the mixers and the residual chlorine content was measured. The mixing qualities of the six bleach lines were compared based on the intensity of segregation. Ventresca et al. (2002) investigated the effects of viscosity ratio of two miscible liquid streams and Reynolds number on the degree of mixing downstream of a static mixer. The two liquids were introduced using a T-junction, and the mixing quality was evaluated by imaging a downstream cross-section of pipe using laser induced fluorescence (LIF). Four statistical methods were utilized to quantify the mixing effectiveness: correlograms, scale of segregation, coefficient of variation (CoV), and intensity histograms. The scale of segregation, representing the clump size, decreased as the flow velocities increased for both low and high viscosity ratios. The CoV, representing the intensity of segregation, also decreased with increasing velocity.  16  2.6 Mixing Assessment on Laboratory and Pilot Scales of Pulp Suspensions Several studies have been conducted to assess mixing quality in pulp suspensions at laboratory and pilot scales. Measurement techniques are based on pulp quality parameters and tracer measurement. Paterson and Kerekes (1985) measured the residual chorine concentration of samples taken from a pulp suspension by a microsampler. The degree of mixing was evaluated in the fibrescale range based on Is and Ls. This method is tedious and time-consuming. Kouppamaki (1985) assessed mixing quality by using a radioactive tracer. A short-lived tracer, Ba-137m, was chosen to reduce risks from radiation. The radioactive tracer did not interfere with process conditions, and this method requires minimal effort since the relative tracer concentration can be measured directly by detectors mounted on the pipe and vessel walls. However, radiation risks are not completely eliminated since the radiation is dangerous, even for a short time. Radioactive generators and detectors are also relatively costly. Breed (1985) evaluated high-shear mixer performance by using sodium chloride as an inert tracer. Pulp samples were collected at the discharge of the mixer and titrated to determine the chloride ion concentration. The NaCl concentration distribution was then determined to assess mixing on a pilot scale. For this study, macroscale mixing was measured. In mill applications, lithium chloride can be used as a tracer instead of sodium chloride since the concentration of sodium ion is relatively high in bleaching processes. Bennington et al. (1997)  17  used lithium chloride as an inert tracer to evaluate mixing quality of four different laboratory mixers over a range of mixing scales. LiCl solution was added to the pulp suspensions before mixing. Sample of various sizes were taken from each mixer, and the mixing quality was evaluated based on the CoV. This method can be applied to any bleaching stages, but it is time-consuming and does not yield continuous (on-line) measurement. Kamal and Bennington (2000) used two fluorescent dyes, mephenesin and 2-naphthalene sulphonic acid, as tracers in mixing studies, with a fibre optic probe measuring the tracer concentration. This method evaluates the mixing quality on a fibre-scale and can be applied as an on-line measurement technique since fluorescent dyes have very fast responses. However, tracers react rapidly with common bleaching chemicals including chlorine dioxide and hydrogen peroxide, making it difficult to evaluate mixing in several bleaching processes. This method can be applied to assess mixing in a stock chest, headbox of a paper machine and a NaOH stage in a bleach plant. Kourunen et al. (2011) applied ERT to assess the mixing efficiency of a pilotscale medium consistency mixer and compared the results with those from temperature measurement. ERT has advantages relative to other techniques since it is non-intrusive, not tedious and able to quantify the mixing throughout the entire suspension volume. In this study, cold water, air and steam were used as tracers, with mixing quality based on the CoV. The softwood pulp suspension mean velocity was ~2.1 m/s at Cm = 10%. ERT clearly indicated worse mixing when the tracer was introduced, and improved mixing when the mixer was turned on, with results similar  18  to those from temperature profiling. However, few trials were performed, and more experimental data are needed.  2.7 Mixing Assessment on the Industrial Scale in Pulp and Paper Processes There have been several attempts to evaluate mixing quality on the industrial scale. The techniques should not interfere with process conditions or adversely affect pulp quality. Most previous techniques are based on extracting samples from the process and analyzing them in the laboratory. Some studies assessed mixing by in-situ measurements, such as a temperature profiling. One of the first attempts to assess mixing quality was by Elliott and Farr (1973), based on chlorine and oxidation reduction potential (ORP) measurements. An ORP probe was inserted through a ball valve located after the chlorine mixer, and ORP curves were determined from the collected data. This technique is simple, without adding any substances to the process. It is, however, intrusive and can be applied only to the chlorine stage of a bleach plant. Torregrossa (1983) assessed mixing quality with LiCl as an inert tracer. The LiCl was added to the ClO2 supply, and small pulp samples were taken at several points downstream of the bleaching tower. The samples were analyzed for consistency and LiCl concentration, and the average ClO2 charges and ClO2 charge deviation were determined. A similar technique was utilized by Backlund et al. (1987), but they used the coefficient of variation as a mixing index. Pulp samples 19  were taken at various positions on the surface of bleaching tower, and the lithium content was measured by flame spectroscopy. Although inert tracer techniques can be applied to any stage in a bleach plant, they are time consuming and not continuous. Paterson and Kerekes (1986) implemented a residual chorine measurement technique to assess mixing quality on an industrial scale. A line sampler obtained samples from the process between the mixer and chlorination tower. While most techniques used to determine mixing quality in a mill provide macroscale measurements, this one evaluates mixing on the fibre-scale. This method is, however, time consuming, not continuous and can be applied only to a chlorine stage. Another method involves measuring temperature around the exterior of the process piping. This method can be utilized when there is a suitable temperature difference between chemical and pulp suspension, as in a ClO2 stage. The ClO2 entering a process at ~10C acts as a temperature tracer in a hot pulp stream at ~70C. Several attempts were made to use this technique to assess mixing in a number of mills (Torregrossa, 1983; Sinn, 1984; Pattyson, 1985; Robitaille, 1987; Rewatkar et at., 2002). Surface temperature was measured at six to eight positions around the mixer discharge piping, with temperature variation utilized to determine mixing quality. Chemical savings could also be estimated by performing mass and energy balances on the ClO2 charge (Rewatkar et al., 2002). This method is nonintrusive and does not interfere with the process. It is also a continuous, simple and efficient technique. However, this technique quantifies mixing only at the pipe  20  periphery, not over the entire pipe cross-section, and a suitable temperature difference between chemical and pulp suspension must exist. Bennington et al. (2001) utilized mixing performance curves obtained from hand-bleaching pulp to evaluate mixing of industrial pulp mixers. Pulp samples were collected before mixers in a bleach plant and used to prepare hand-bleaching pulp in the laboratory. The kappa number (representing lignin content) was measured for various chemical charges to determine the mixing performance curves for various mixing conditions, based on the coefficient of variation. Time-tracked samples of pulp were collected downstream of bleaching towers based on predicted tower residence times. The kappa number was then measured and used to evaluate mixing quality based on mixing performance curves. This method has the advantage of using pulp itself as a tracer. However, the measurement is time consuming and not continuous since pulp samples need to be taken from the process.  2.8 Electrical Resistance Tomography (ERT) 2.8.1 Principles and system structure Electrical resistance tomography (ERT) is a non-intrusive technique that can be used to determine the distribution of electrical conductivity in process vessels and pipelines from measurements at the vessel periphery. ERT collects data based on the adjacent strategy (Dickin and Wang, 1996). Electrical current is applied  21  through a pair of neighbouring electrodes, and voltage differences are measured between remaining pairs of neighboring electrodes. Current is then applied through the next pair of electrodes and the voltage measurements are repeated. The measurement process is repeated for all other pairs of current-injecting electrodes, and, with the aid of suitable commercial software for image reconstruction, the spatial gradients of electrical conductivity in a cross-section can be determined. ERT can be used effectively when the main continuous phase in the process is at least slightly conductive and the other phases have different conductivities. ERT provides excellent temporal resolution, but achieves limited spatial resolution. This level of spatial resolution is, however, adequate for many industrial applications, and, importantly, the high temporal resolution provides fast real-time imaging (York, 2001). The spatial resolution depends on the number of electrodes, the noise in the measurements, the conductivity difference between phases and the position in the image (Seagar et al., 1987). It also varies with the image reconstruction algorithm. Holden et al. (1998) tested the spatial resolution of an ERT system with 16 electrodes per plane. The data collection and image reconstruction were based on the adjacent strategy and a linear back projection method, respectively. Two different sizes of aluminum rods and a PVC rod were inserted into different regions in a cylindrical vessel equipped with the ERT system. The conductivity regions in the reconstructed image were compared to the locations of the rods, and the spatial resolution was estimated to be of the order of 4.2% of the vessel diameter.  22  The ERT system structure consists of three main components: sensor system, data acquisition system (DAS) and image reconstruction system (Dickin and Wang, 1996).  2.8.1.1 Sensor system In order to obtain reliable measurement, the electrode material should be more conductive than the fluids being imaged. Metallic electrodes, e.g. stainless steel, silver, gold, platinum or silver/palladium, are often utilized for process applications (Dickin and Wang, 1996). The number of electrodes (ne) affects the sampling time and the spatial resolution. The time for data acquisition and image reconstruction is a function of ne, and the spatial resolution is proportional to N m , where Nm is the number of independent measurements (Dickin and Wang, 1995). For the adjacent strategy, the total number of measurements is ne2, but only ne(ne 1)/2 are independent. In addition, the voltage differences are not measured at a pair  of  current-injecting  electrodes,  to  avoid  electrode/electrolyte  contact  impedance problems, reducing the number of independent measurements (Nm) to ne(ne - 3)/2 (Dickin and Wang, 1996). The number of electrodes is typically 8, 16 or 32, with a maximum of 32 to maintain optimal flexibility and accuracy (Dickin and Wang, 1995). The electrodes should be positioned at equal intervals to obtain the maximum amount of information from the inside of the vessel. The most common electrode geometry for pipelines is a set of 16 electrodes spaced at equal intervals around the boundary of a circular vessel. 23  The size of the electrodes is another important factor. To obtain an even current density within the vessel, a large surface area is required for the currentinjecting electrodes, whereas a small surface area is required for optimal voltage measurement (Dickin and Wang, 1996). Wang et al. (1995) showed that errors in measurement from using larger electrodes were negligible. Thus, utilizing the same electrode size for current injection and voltage measurement is generally accepted. The amplitude and frequency of Injected current can be adjusted to optimize the signal-to-noise ratio of the measured voltage signal. For slowly changing processes, more accurate measurements are obtained at lower frequencies of injected current (Dickin and Wang, 1995). A relatively low excitation frequency, e.g. 9.6 kHz, is generally applied for ERT (Mann et al., 1997).  2.8.1.2 Data acquisition system The DAS, mainly consisting of a signal source, an electrode multiplexer array, voltmeters, signal demodulators and a system controller, obtains the quantitative data for image reconstruction (Mann et al., 1997). Data collection must be fast and accurate in order to track small changes of conductivity in real-time, resulting in precise measurement of the conductivity distribution. The signal source consists of a master oscillator and a voltage-to-current converter (Plaskowski et al., 1995). The oscillator generates a sine-shaped voltage signal fed into a voltage-tocurrent converter, referred to as a voltage-controlled current source (VCCS). Multiplexers are used to share the current source and voltage measurement stages between any numbers of electrodes (Dickin and Wang, 1996). Finally, signal 24  demodulators are utilized to recover the amplitude attenuation and phase shift of the sine wave signal as a result of passing through a resistive medium in order to optimize the signal-to-noise ratio.  2.8.1.3 Image reconstruction system The image reconstruction process determines the internal distribution of electrical conductivity within the process vessel based on two reconstruction algorithms. The first algorithm solves the “forward problem”, by creating images depicting a change in conductivity relative to a set of reference data acquired initially, whereas the other solves the “inverse problem” by producing images depicting the conductivity of each pixel (Dickin and Wang, 1996). It is necessary to solve the forward problem since the measurements taken at the boundary of process vessel contain insufficient information to solve the inverse problem directly. Finite element methods are widely used to solve the forward problem, by determining the electrical potential and potential differences between electrode pairs for a given current injection from the known distribution of conductivity, i.e. the image, to obtain the sensitivity matrix (Giguere et al., 2008). Based on the sensitivity matrix, the inverse problem can be solved to determine the distribution of conductivity by using an appropriate quantitative algorithm. Several algorithms have been utilized, including linear back projection, Landweber, Newton-Raphson, Tikhonov and Levenburg-Marquardt, with details provided by Giguere et al. (2008). The non-iterative linear back projection is simple and fast, allowing on-line imaging, whereas the iterative methods typically provide 25  more accurate images, but are more time-consuming and may lead to problems with convergence (York, 2001). The commercial software utilized in the present project from Industrial Tomography Systems, Manchester, UK solves the forward problem by the finite-element method, and utilizes the linear back projection algorithm to solve the inverse problem.  2.8.2 ERT applications on the industrial scale ERT is an inexpensive and fast visualization tool for various applications on the industrial scale. Williams et al. (1999) utilized ERT to monitor and improve industrial hydrocyclone performance through discharge fault detection, auditing of operations and quantification of radial concentration profiles. Eight planes with 16 electrodes per plane were located on the internal surface of the hydrocyclone, ranging from the feed to the outlet, and a linear back projection scheme was employed for image reconstruction. The location of the air core characterized by a region of low conductivity was visualized for different fault conditions. The size of the air core was also estimated as a function of pressure drop. The dynamics of this process were so fast that rapid data acquisition was required (York, 2001). Vlaev et al. (2000) applied ERT to investigate three aspects of filter monitoring: liquid level during drainage, tilt of filter assembly and malformation in a wet cake. Due to the conductivity difference between the filtrate liquor and filter-cake, ERT was able to detect malformation of the wet cake. ERT was also able to measure conductivity changes accompanying the fall of liquor level above the filter cake and plate. For this process, the dynamics were slow and modest data acquisition, of the order of 26  one frame per minute, was adequate (York, 2001). Henningsson et al. (2007) implemented ERT to monitor the rinsing step in a dairy plant. The displacement of yoghurt filled in the pipe by water flow was visualized, and temporal variation of yoghurt concentrations at various positions in the cross-section was determined, assuming a linear dependency on conductivity. The ERT results indicated lower yoghurt concentration at the centre of the pipe than at the top and bottom, consistent with the CFD simulation results.  2.8.3 Use of ERT for quantification of mixing Application of ERT to study mixing processes provides better understanding of the processes and optimization of mixer design and operation. The degree of mixing can be evaluated both qualitatively by imaging the processes and quantitatively by using mixing indices based on statistical analysis of image pixel values. A number of ERT applications to mixing processes have been reported. Williams et al. (1996) used ERT to determine the effect of impeller type, agitation speed and particle size on solid-liquid mixing. Four planes with sixteen electrodes per plane were utilized to evaluate the mixing quality of sand/water mixtures in a stirred vessel. The solid distribution was quantified by the CoV based on the concentration in each image pixel for each measurement plane. Due to the presence of a central metal impeller, with higher conductivity than the bulk slurry, the average conductivity of a given plane was determined from the outer 48 pixels of the finite element mesh for the lower two planes. The overall mixing index for the reactor volume was evaluated by averaging the mixing indices for the four planes. 27  Wang et al. (2000) measured gas-liquid mixing in a stirred vessel using ERT. An 8-plane 16-ring-element ERT system was applied to determine air-water lowviscosity mixing and air-carbopol high-viscosity mixing. The ERT system detected and reconstructed 2D images caused by local variation of dispersed gas holdup or voidage. 3D images were also created by interpolation between the planes. Kaminoyama et al. (2005) investigated dispersion of liquid monomer and polymer droplets in a stirred-tank polymerization reactor with a six-blade Rushton turbine impeller. The influences of impeller rotational speed and the injection amount of dispersing agents on the dispersion state were quantified by a dispersion index, derived based on the CoV of ERT image pixel values. Kim et al. (2006) applied ERT to monitor the mixing of two miscible liquids in a cylindrical stirred vessel with a six-blade Rushton turbine impeller, based on quantification of image pixel values. The temporal variation of the CoV of crosssectional conductivities was evaluated to determine the mixing time, the time taken to reach a uniform concentration distribution. Stephenson et al. (2007) utilized ERT to determine jet mixing in water flow for both coaxial and side entry jets. The experiments were conducted: (a) to visualize the plume of tracer formed by jets adjacent to the jet entry and (b) to monitor mixing along the length of the pipe. Sodium chloride was added in the jet addition stream as a tracer for a three-plane of 16 electrodes per plane sensor array. The CoV was utilized as a mixing index. The overall mixing index and mixing length were determined from the exponential decay of conductivity variance defined as:  28    2 ( z)  ln 2     z   (0)   (2.1)  where z is the distance downstream of the tracer addition,  2 (0) the conductivity variance before tracer addition and  the overall mixing index. The results showed that ERT is an appropriate tool for studying jet mixing in pipelines, both for visualizing the tracer plume formed by jets adjacent to the jet entry and for monitoring mixing along the pipe. Hosseini et al. (2010) examined the influences of impeller type, impeller speed, off-bottom clearance, particle size and concentration on solid-liquid mixing in a stirred tank based on ERT. The degree of homogeneity in the mixing vessel was determined from the standard deviation of the averaged solid concentration for each of eight sensor planes. The impeller speed was set to a desired value and reference measurements were taken at steady state to eliminate effects of impeller and system noise. For several types of impeller, mixing improved with increasing impeller speed, but deteriorated with further increase in impeller speed once the homogeneity reached a maximum. Optimum design parameters and operating conditions were also developed. Tahvildarian et al. (2011) applied the same technique to evaluate mixing of micron-sized polymeric particles in a slurry reactor. Mixing deteriorated with increasing particle size and concentration. Optimum impeller design parameters including impeller type, speed and diameter were also established.  29  2.8.4 Use of ERT in pulp and paper systems ERT has been utilized extensively in pulp and paper systems including model chip digesters (Vlaev and Bennington, 2004; Vlaev and Bennington, 2005; Vlaev and Bennington, 2006; Ruzinsky and Bennington, 2007; Lee and Bennington, 2007; Lee and Bennington, 2010), agitated pulp stock chests (Hui et al., 2009; Bhole and Bennington, 2010; Bhole et al., 2011), a semi-batch slurry bubble column (Ishkintana and Bennington, 2010) and a medium consistency mixer (Kourunen et at., 2011). Vlaev and Bennington (2004) visualized liquor flow through a model digester using eight ERT sensor planes with sixteen sensors per plane spaced equally around the vessel periphery. An ITS P2000 system provided data acquisition, with image reconstruction based on a linear back projection algorithm. A brine solution, with a different conductivity from the reference conductivity, acted as a tracer in the experiments, injected at two locations: through the downcomer and into the flow entering the base of digester. The ERT results were similar to optical images for various flow situations, and agreed well with industrial field observations. Later, they extended the study to determine the optimum flow ratio for creating a uniform zone within the digester (Vlaev and Bennington, 2005). Two indices were utilized: the mutual bulk resistance (MBR) of a sample plane and the CoV of the image pixels. The former is to determine when steady-state conditions are established and the latter is to evaluate the relative uniformity as a function of time. Ruzinsky and Bennington (2007) applied ERT to determine the effect of the particle to vessel size ratio on liquor flow uniformity, comparing the results of a packed column of wood chips with smaller HDPE pellets. The flows adjacent to the 30  digester wall and exiting the central downcomer were also examined. A brine solution was used as a tracer, with a step change in tracer concentration in the fluid entering the downcomer causing a drop in conductivity from 1.0 to 0.78 mS/cm. Lee and Bennington (2007) measured flow uniformity in a model batch digester based on ERT. A brine solution again served as tracer. The electrical conductivity of the fluid in the vessel was set at 3.0 mS/cm at the beginning of the test. Fluid of conductivity 2.5 mS/cm was then introduced as a step change, with ERT imaging the lower conductivity tracer. Local flow velocities were also determined based on pixel-pixel cross-correlation techniques. Lee and Bennington (2010) compared experimental ERT results with computational predictions from a commercial CFD code, assuming the model digester to be uniformly packed by incompressible solids with a uniform void fraction. Hui et al. (2009) utilized ERT to measure cavern size and shape in pulp suspensions in a cylindrical stock chest equipped with a side-entering Maxflow impeller for various operating conditions, including impeller speed, impeller offset from the wall and suspension height, with 1.0 – 5.0%wt. softwood and hardwood pulp suspensions. Conductive tracer particles were used to visualize caverns instead of saline solution, to avoid concerns about potential tracer diffusion overestimating the cavern size. The cavern area for each measurement plane and the total cavern volume measured by ERT were similar to the results obtained from ultrasonic Doppler velocimetry, with considerably less time required by ERT for data acquisition and analysis. A model for estimating cavern size that includes cavern-  31  wall interaction was also developed, with predicted cavern volume agreeing relatively well with the experimental results. Bhole and Bennington (2010) measured cavern size in 3% by wt. hardwood pulp suspension for four axial-flow impellers based on ERT, and compared the results with predictions of the model developed by Hui et al. (2009). Bhole et al. (2011) extended the previous work by Hui et al. (2009) to investigate the influence of off-wall clearance of the side-entering Maxflow impeller on cavern size in hardwood pulp suspensions with fibre mass concentration of 2 – 4%, and compared the cavern sizes measured by ERT and predicted by CFD simulations with predictions of the model of Hui et al. (2009). Ishkintana and Bennington (2010) investigated gas hydrodynamics in pulp suspensions in a semi-batch bubble column using ERT. Local gas holdups were determined from local conductivity values based on the Maxwell equation. Salt was added to the water to increase the background conductivity, maximizing the signalto-noise ratio and improving image reconstruction. Axial and radial gas holdup profiles were also examined. The average gas holdups obtained from ERT agreed well with those from other techniques, including height difference and pressure difference methods, for various superficial gas velocities. Recently, Kourunen et al. (2011) measured the mixing efficiency of a pilot-scale medium consistency mixer based on ERT and CoV of local conductivity, with details discussed in section 2.6. They showed that ERT was efficient in assessing both liquid-suspension and gassuspension mixing, with high potential for implementing as a real-time monitoring tool in industrial bleaching processes.  32  In summary, ERT has been applied successfully to a number of pulp and paper systems. With the use of multi-plane electrical sensors, ERT can provide continuous assessment of mixing of pulp suspensions. Previous studies show that process flows can be imaged effectively when the conductivity difference between the background and tracer solution is at least 20%. Since the conductivity difference between the pulp suspension and the added chemical is ~60% in the D0 stage, as shown in Table 2.1, there is potential to apply ERT to image industrial pulp suspension flow and to measure mixing quality in the first chlorine dioxide bleaching (D0) stage at Howe Sound Pulp and Paper. However, ERT has limitations that need to be recognized from the outset: (a) spatial resolution of only 5 – 10% of the pipe diameter; (b) need for a custom sensor array to be fabricated and installed; (c) requirement that there be a significant difference in electrical conductivity between the two fluids being mixed.  33  3. Experimental Details and Mixing Indices 3.1 Experimental Apparatus2 Experiments were conducted in a pilot-scale flow loop facility, consisting of two 4000-l tanks, a 40-HP Allis Chalmers PWO centrifugal pump equipped with a variable frequency drive, Wika S-11 gage pressure transmitters upstream and downstream of the pump, a Rosemount 8712 magnetic flow meter with a GF EA31 computer-actuated valve downstream, and a 40-m length of 76.2-mm I.D. PVC pipe. A 10-HP MixPro agitator, model 20ECO (Professional Mixing Equipment Inc., Brampton, ON), powered by an electric motor was installed to agitate the pulp suspension to maintain homogeneity in the tank. The acrylic test section was fabricated and installed 5 m downstream of the closest pipe bend, ensuring that the flow was fully developed before entering the test section. A schematic diagram of the flow loop is presented in Figure 3.1. The acrylic test section is 76.2 mm in diameter and 1.88 m long. Eight sensor planes are spaced at 250 mm intervals along the pipe, with the first of these planes 67 mm upstream of the 90 side-stream injection, where liquid or gas is introduced  2  Part of this section is similar to part of a paper that has been published: Yenjaichon, W., Grace,  J.R., Lim, C.J., Bennington, C.P.J., 2012. In-line jet mixing of liquid-pulp-fibre suspensions: Effect of concentration and velocities. Chem. Eng. Sci. 75, 167–176, and a paper accepted for publication: Yenjaichon, W., Grace, J.R., Lim, C.J., Bennington, C.P.J., 2012. Pilot-scale examination of mixing liquid into pulp fiber suspensions in the presence of an in-line mechanical mixer, Ind. Eng. Chem. Res.  34  into flowing water or pulp suspensions. Each sensor plane contains 16 stainless steel electrodes of diameter 6.35 mm, flush with the inside wall of the test section and spaced at regular 22.5 intervals around the pipe periphery. A ground electrode of the same size as the sensor electrode is located between the P2 and P3 planes. The internal diameter of the side injection tube, Dj, varied from 3.81 to 15.9 mm, corresponding to jet-to-pipe diameter ratio, Dr, of 0.05 to 0.208, covering literature and industrial diameter ratios. The injection tube was flush with the inside wall of the test section, with sufficient length to provide fully developed flow before entering the main stream. Two SSI MediaGauge pressure gauges were installed 100 mm upstream and 230 mm downstream of the test section. A schematic of the test section appears in Figure 3.2.  A  B  Figure 3.1: Schematic of pilot-scale flow loop facility. 35  Dj  Figure 3.2: Schematic of test section (all dimensions in mm).  As suggested by the ERT instrument supplier, the electrode size is equal to the gap size (Bolton, 2008). For the present study, an electrode width of 7.5 mm is recommended for the 76.2-mm I.D. test section. However, due to the availability of materials, the electrode size of 6.35 mm (¼ inch) has been used for the sensor system. A large surface area of the current-injecting electrodes is required when dealing with high conductivity fluids to ensure that an even current density is generated within the vessel (Dickin and Wang, 1996). Since the main fluids in this study are water and pulp suspensions, with relatively low conductivity, a slight decrease in electrode size likely has insignificant effect on the accuracy of the measurement. A commercial electrical resistance tomography system, model P2000 (Industrial Tomography Systems, Manchester, UK), is utilized for data acquisition. The electrodes are connected to the ITS P2000 system via 2 m long coaxial cables. The test section and ERT system are illustrated in Figure 3.3.  36  Ground electrode  P2000 Unit  P3 Test Section  P8  Figure 3.3: Photo showing test section and ERT system.  The ERT system applies a constant AC current to a pair of electrodes and measures the voltage differences between the other electrode pairs using an adjacent-pair strategy (ITS, 2007) with an excitation frequency of 9.6 kHz and an injection current of 15 mA. Two versions of linear back-projection algorithm were employed for image reconstruction including the Modified Sensitive Back Projection (MSBP) for liquid-liquid/suspension mixing and the Sensitive Back Projection (SBP) for gas-liquid/suspension mixing. The latter provides more representative images for the systems with high contrast in conductivity such as air-water system (ITS, 2007). The sampling time interval was 55 ms, with four samples averaged. Details of the parameter setup are provided in Figure 3.4.  37  Figure 3.4: Parameter setup for electrical resistance tomography. In our case, ERT achieves a spatial resolution of 5 - 10% of the pipe diameter, i.e. 3.8 to 7.6 mm for the 76.2-mm ID test section, representing the upper size of the fibre-scale. The scale of measurement is the macroscale at the pulp mill, with radial measurement increments from 30.5 to 61 mm, based on the pipe ID of 610 mm for the first chlorine dioxide (D0) stage at Howe Sound Pulp and Paper (HSPP). For an axial resolution, the macroscale is measured. In the laboratory, an ITS P2000 was utilized at a rate of data acquisition of ~1 Hz, whereas an ITS Z8000 (Industrial Tomography Systems, Manchester, UK) was used at the D0 stage at HSPP, with data acquired at a rate of ~10 Hz. The mainstream velocity varied from 38  0.5 to 5.0 m/s on the laboratory scale, and from 0.8 to 1.4 m/s for the industrialscale tests, with the axial measurement increments of 50 to 500 cm and 80 to 140 mm, respectively. For some flow patterns in horizontal gas-liquid flow such as stratified, elongated bubble and slug flow, some electrodes lose electrical contact, leading to anomalous voltage measurements. Resistor adaptors obtained from the ERT system supplier and inserted between electrodes and the ITS P2000 unit, were utilized for all gas-liquid/suspension tests (Chapters 6 and 7) to provide a by-pass route for the electrical current, improving reconstructed images for these flow conditions. The liquid injection system consists of a 120-l tank, a 1.5-HP pump, model JL3513A (Baldor Electric Co., Fort Smith, AR) and a Rosemount 8712 magnetic flow meter. The flow rate is controlled manually by adjusting the valve opening. The injection system was connected to an injection tube at the bottom of the test section via a 6.8-m length of 12.7-mm I.D. PVC flexible tube and a 25.4-mm I.D. brass spring-loaded piston check valve. The gas injection system, including an air filter/regulator, a Parker Domnick Hunter prefilter (model 00261JN), a Cole-Parmer mass flow meter (model 32908-77), a brass adjustable pressure relief valve and 4.7-m length of 12.7-mm I.D. copper tube, was fabricated for the gas mixing experiments which are the subjects of Chapters 6 and 7. The system was connected to the injection tube at the bottom of the test section via a 2.8-m length of 12.7-mm I.D. EPDM rubber hose and a 25.4-mm I.D. brass ball valve.  39  In some cases, a Phantom v611 high-speed camera (Vision Research, Inc., Wayne, NJ) was used from the side to capture images of gas-liquid flow along the pipe at 1,000 frames/s and a shutter speed of 10 s, providing an image size of 800 × 600 pixels (spatial resolution of ~0.18 mm per pixel). Images and flow patterns for different flow conditions were then compared with the ERT results. The images are not of a quality suitable for inclusion in the thesis, but are sufficiently clear to identify the flow patterns. The mechanical mixer system consisted of a ½ HP motor (LEESON Electric, USA), a 1.41 N-m rotating torque sensor with integral optical encoder (OMEGA Engineering, Inc., USA), a 12.7 mm diameter stainless steel shaft and a 63.5 mm wide octagonal impeller with a 38.1 mm diameter hole in the middle to prevent excessive force on the impeller when it was perpendicular to the flow. The octagon minimizes the gap between the impeller and the internal wall of the pipe, hence maximizing the shear. This type of impeller was chosen by the late Dr. Bennington who initiated this work, presumably to be simple and to have some similarity to the high-shear mixers. A schematic diagram of the test section and in-line mechanical mixer is shown in Figure 3.5. The torque from the impeller was not measured since the frictional torque losses at the shaft seals exceeded the impeller torque and depended on the condition of the seals. The losses increased substantially as small fibres entered the seal gaps, and decreased significantly when the seals became worn.  40  Pipe  Impeller  Shaft  (6.35 mm thick)  Hole  63.5  6.35 12.7 25.4 38.1 63.5  6.35  76.2  Figure 3.5: Schematic of in-line mechanical mixer (all dimensions in mm).  41  3.2 Materials and Experimental Methods Two pulp types, one softwood and the other hardwood kraft, were tested in this study. The pulp sheets used to prepare pulp suspensions were Northern Bleached Softwood Kraft (NBSK) pulp from Canfor Pulp Ltd., Prince George, BC, with fibres of arithmetic average diameter 27.6 m and coarseness 0.129 mg/m, and length-weighted average length 2.52 mm. The hardwood pulp was bleached maple kraft pulp from Domtar Corp., Espanola, ON, with fibres 0.60 mm long, 16.6 m in diameter and 0.066 mg/m in coarseness on average. Coarseness is defined as the weight per unit length of fibres. The fibre properties were determined by a Fibre Quality Analyzer (FQA) (OpTest Equipment Inc., Hawkesbury, ON). The length-weighted average is utilized since it reduces impacts from fines on the length measurement (OpTest Equipment Inc., 2008). Pulp suspension tests were conducted at four fibre mass concentrations (also called “consistency”) (0.5, 1.0, 2.0 and 3.0%), defined as weight of fibres per weight of suspension, expressed in wt.%. Pulp sheets were prepared according to the amount of pulp suspension in the tank (3,200 – 3,800 L). The sheets were torn by hand and diluted with tap water to obtain a desired fibre mass concentration. The fibres were dispersed and separated from one another by an agitator in the tank. The agitator was operated at a high enough speed (~100 rpm for water, up to 350 rpm for pulp suspensions) to ensure the homogeneity of pulp suspension in the tank, visually observed by vigorous motion of pulp suspension throughout the entire tank. Three samples were taken from the tank for consistency measurements. Sodium chloride salt was added to the tank to increase the background electrical 42  conductivity to ~0.22 mS/cm (equivalent to a salt concentration of ~0.1 g/L), to improve the accuracy of the ERT measurements. Steady state experiments were conducted with constant mainstream and side-stream flow rates and (where applicable) impeller rotational speed, with the fluid temperature in the main stream maintained at 20 ± 5C. The mainstream flow rate was controlled by adjusting the pump frequency or the opening of the computer-actuated valve downstream of the flow meter. Prior to collecting ERT data, a reference measurement was taken with only the main stream present (i.e. before side-stream injection) as a basis for subsequent measurements.  3.2.1 Test procedures for liquid injection In Chapters 4 and 5, water or pulp suspension from tank A (or B) (see Figure 3.1) was pumped through the loop at a constant flow rate, and a brine solution, with concentration from 0.9 to 2.4 g/L, then entered at a 90 T-junction at the inlet of the test section. The mixture flowed through the test section and back to tank A (or B), to maintain the fluid level in the tank and hence the mainstream flow rate. This led to an increase in salt concentration over time, but the effect was negligible due to large volume of fluid in the tank (3,200 – 3,800 L) and brief duration of each test (~70 s): the salt concentration increased by only ~4% for the highest amount of injected tracer and the highest fibre mass concentration (hence the lowest amount of water in the tank) tested. The tracer concentration was adjusted after each test by adding salt into the side stream to match the salt concentration in the main stream, maintaining a jet-to-pipe NaCl concentration ratio of ~10. Since it is more 43  convenient to measure the electrical conductivity during the tests, and the electrical conductivity is directly proportional to the salt concentration, as shown in Figure 3.6, the ratio of conductivity in the jet stream to that in the main stream was kept at ~10. For all the tests in this study, the temperature difference between the main and side stream was less than 10C.  Electrical conductivity (mS/cm)  6  Y = 2.124X 2 R = 0.9994  4  2  0 0  1  2  3  NaCl concentration (g/L)  Figure 3.6: Electrical conductivity as a function of NaCl concentration.  3.2.2 Test procedures for gas injection In Chapters 6 and 7, air was introduced into the flowing water or pulp suspension at a 90 tee at the inlet of the test section. The developing flow conditions were investigated in the entrance region downstream of the tee. The air flow was determined based on the pressure and temperature at the injection point 44  with the inlet pressure measured by a pressure gauge at the inlet of the test section, and the temperature determined from a sample taken from the tank. Ideal gas behaviour was assumed in correcting the air flow rate set point at injection conditions to standard conditions (101.3 kPa and 25 C), at which the readings were shown in the flow meter. The superficial liquid/pulp velocity and superficial gas velocity were adjusted by changing the pump frequency and the gas injection system control valve, respectively.  3.3  Data Analysis Image reconstruction was performed by the commercial ERT software to  calculate the cross-sectional distribution of electrical conductivity within each measurement plane. An image contains the conductivity values in a square grid with 20 x 20 = 400 pixels. For the circular pipe, the image is reconstructed using 316 pixels from the 400 pixel square grid, excluding pixels lying outside the pipe periphery, as shown in Figure 3.7. A tomographic image shows the conductivity distribution, with low-conductivity regions blue and high conductivity regions red. A sample is illustrated in Figure 3.8. Due to misalignment between the electrode position in the ERT reconstruction process and the actual electrode position in each sensor plane, the top position of the pipe in the tomographic images in Chapters 4 – 7 (laboratory-scale experiments) were rotated counterclockwise by 11 from the vertical axis, whereas the images in Chapter 8 (industrial-scale experiments) were rotated counterclockwise by 90 from the vertical axis. 45  Figure 3.7: ERT image reconstruction grid for pipe circular cross-section (ITS, 2007).  Figure 3.8: Tomographic image showing regions of high and low conductivity (ITS, 2007). Bar at the bottom shows colours corresponding to different conductivities. 46  3.3.1 Characterization of liquid-suspension mixing3 The ERT measurement data were analyzed in MATLAB 7.0 (MathWorks, Inc., Natick, MA). The degree of mixing was determined from the measurement data in each cross-sectional sensor plane along the pipeline using the coefficient of variation (CoV) of the individual conductivity values in each image pixel to define as a mixing index  n  (y Mm    y  i 1    mc ,i   y) 2  n 1 y  (3.1)  where  is the standard deviation of the conductivity values, ymc,i is the local mixture conductivity determined from ERT measurements, y is the average conductivity, and n is the total number of pixels in the measurement plane (316 pixels). To account for differences in conductivity due to the inherent heterogeneity of the pulp suspension and system noise, the index defined by equation 3.1 was corrected by  M  M m2  M s2  (3.2)  where Ms is the system mixing index measured at P1 (just upstream of tracer injection), averaged over 50 values obtained during the same measurement intervals  3  A version of this section has been published: Yenjaichon, W., Grace, J.R., Lim, C.J., Bennington,  C.P.J., 2012. In-line jet mixing of liquid-pulp-fibre suspensions: Effect of concentration and velocities. Chem. Eng. Sci. 75, 167–176.  47  as those in which data were obtained for planes P2 to P8. In addition, in order to exclude the effect of differing amounts of tracer for different flow conditions and to adjust the mixing index to a scale of 0-100%, the mixing index is normalized with CoV for fully segregated flow (CoVFS or MFS), defined as:  CoV FS  M FS  n  where  Thus  and  2  FS    (C i 1  i   C)2  n    Qp Q j  Qp  M FS   M    FS  (C p  C ) 2   Q pQ j Q jC j  QpC p  M  M FS  (3.3)  C  Qj Q j  Qp  Cp  Cj  M m2  M S2 M FS  (C j  C ) 2  (3.4)  (3.5)  (3.6)  M’ is the modified mixing index, FS the standard deviation for fully segregated flow, Ci the local salt concentration in the mixture, C the average salt concentration of the mixture, Cp the salt concentration in the main stream or pipe flow, Cj the salt concentration in the side (jet) stream, Qp the volumetric flow rate of the main stream, and Qj the volumetric flow rate of the side jet stream. Fully segregated flow would occur if the mainstream and side-stream flows were immiscible or totally unmixed (with each sample composed solely of one stream or the other), the worst possible mixing condition for a specific amount of tracer, mainstream flow and sidestream flow. The modified mixing index decreases as the mixing quality improves 48  and reaches zero for perfect mixing. On the other hand, it is 1.0 (i.e. 100%) for fully segregated flow. This index is analogous to the intermittency index characterizing density fluctuations in circulating fluidized beds (Brereton and Grace, 1993). Since the electrical conductivity is directly proportional to the salt concentration as noted above, the modified mixing index based on concentration is equal to that based on electrical conductivity.  3.3.2 Characterization of gas-suspension mixing4 The local gas volume fraction in each cross-sectional sensor plane along the pipe was determined from the local conductivity in each image pixel derived from a Maxwell equation (ITS, 2007):   g ,i   2 y1  y 2  2 y mc ,i  y mc ,i  y mc ,i y 2 y1  y  2 y mc ,i  2 y1  y 2  y1  (3.7)  where y1 and y2 are the conductivities of the continuous and dispersed phases, respectively, and ymc,i is the local electrical conductivity from ERT measurements. Since the gaseous phase is non-conductive (y2 = 0), equation 3.7 simplifies to   g ,i   2 y1  2 y mc ,i y mc ,i  2 y1  (3.8)  4  This section is very similar to part of a paper accepted for publication: Yenjaichon, W., Grace, J.R.,  Lim, C.J., Bennington, C.P.J., 2012. Characterisation of gas mixing in water and pulp-suspension flow based on electrical resistance tomography, Chem. Eng. J.  49  The conductivity of the continuous phase is measured independently using a conductivity meter. The local gas holdup varies from 0 to 1, so that y mc,i = y1 for the liquid phase, and y mc ,i = 0 for the gas phase. The ERT measurement data were again analyzed in MATLAB 7.0 (MathWorks, Inc., Natick, MA). A mixing index for gas dispersion in the main stream was determined from the standard deviation of local gas holdup in each image pixel in a cross-sectional sensor plane:  n  m    ( i 1  g ,i    g )2  n 1  (3.9)  where  g ,i is the local gas holdup obtained from ERT measurements,  g is the average gas holdup in the cross-sectional plane, and n is the total number of pixels in the measurement plane (316 here). Due to the inherent heterogeneity of the pulp suspension and system noise, the mixing index defined by equation 3.9 was modified by   g   m2   s2  (3.10)  where s is the standard deviation measured in the absence of gas at plane P1. The mixing index is then normalized by the standard deviation for fullysegregated flow (g,FS) to eliminate the influence of differing amounts of gas for different flow conditions and to adjust the mixing index scale to 0-100%. Fullysegregated flow is the worst possible mixing condition for a specific amount of gas in the flow.  50  n   g2, FS  Thus   ( i 1  g ,i    g )2  n    g (1   g ) 2  (1   g )(0   g ) 2   g , FS   g (1   g )  n   ( and  g  Mg   g , FS  i 1  g ,i    g )2  n 1  g (1   g )  (3.11)    S2  (3.12)  where Mg is the gas mixing index. The gas mixing index depends on the bubble to image pixel size ratio. Bubbles which are much smaller than the pixels are likely to produce uniform dispersion of gas in each pixel, so that Mg approaches 0. Mg increases when the gas dispersion becomes less uniform for large bubbles or slugs, and approaches 1 (i.e. 100%) for fully stratified flow. This index is again analogous to the intermittency index characterizing density fluctuations in circulating fluidized beds (Brereton and Grace, 1993), and very similar to the “intensity of segregation” introduced by Danckwerts (1952). Mg might more properly be called the gas segregation index rather than the gas mixing index as it varies from 100% for perfect segregation to 0 for perfect mixing (no segregation). However, to be consistent with the liquid mixing indices used, not only in this thesis, but also by other authors in this field, Mg has been defined and labelled in this manner. This also facilitates logarithmic ordinate axes for both liquid and gas mixing. The distribution of gas in a cross-sectional plane was examined based on the radial gas holdup profiles, determined from the local gas holdup along a vertical line bisecting the plane. Due to misalignment between the electrode position in the ERT  51  reconstruction process and the actual electrode position in each sensor plane, the vertical line is rotated counterclockwise by 11 from the vertical axis, and a corresponding corrective rotation was applied when analyzing all tomographic images. In Chapter 6, a measure of relative size of gaseous entities is estimated based on the scale of segregation, Ls (Danckwerts, 1952). A central square of width 53.3 mm (each side containing 14 × 14 = 196 pixels, the largest square that can be produced from the total of 316 pixels in a circular plane) for each cross-sectional plane is used to determine Ls. The coefficient of correlation, R(r), is computed from  R(r )   ( g ,i   g )( g ,i  r   g ) ( g ,i   g ) 2  (3.13)  The area under a plot of R(r) versus r is then used to estimate Ls: r0  Ls   R(r )dr  (3.14)  0  where r0 is the value of r at which R(r) reaches zero. The scale of segregation does not precisely match the average diameter of bubbles or slugs, but its magnitude varies in the same sense as the size of clumps (Danckwerts, 1952). Ls is therefore a useful indicator of the relative size of gaseous entities.  52  4. In-Line Jet Mixing of Liquid-Pulp-Fibre Suspensions5 4.1 Introduction Chemical injection is commonly used as a pre-distributor to mix liquid chemicals into pulp suspensions ahead of or inside various mixers, such as static mixers and high-shear mixers, with details provided in section 2.2. The simplest type is a tee mixer. Understanding in-line jet mixing behaviour is essential since it provides basic concepts and guidance for the design of in-line mixers and for process optimization. One of the most popular approaches for in-line mixing is the use of side-entry jet mixers. Good mixing can be achieved when the ratio of the momentum of the inlet jet to the momentum of the main stream is sufficiently high to blend the side stream with the main stream. If the jet momentum is too low, it is insufficient to allow the side stream to escape from the boundary layer at the near-pipe-wall, and the side stream becomes sidewall injection. On the other hand, if the jet momentum is too high, the side stream impinges on the far wall of the pipe, and the impingement point moves closer to the inlet as the jet momentum increases. This condition might not be desirable from a practical point of view since it creates significant stress on  5  This chapter is very similar to parts of a paper that has been published: Yenjaichon, W., Grace,  J.R., Lim, C.J., Bennington, C.P.J., 2012. In-line jet mixing of liquid-pulp-fibre suspensions: Effect of concentration and velocities. Chem. Eng. Sci. 75, 167–176, and a paper accepted for publication: Yenjaichon, W., Grace, J.R., Lim, C.J., Bennington, C.P.J., 2012. In-line jet mixing of liquid-pulp-fiber suspensions: Effect of fiber properties, flow regime and jet penetration, AIChE J.  53  the opposite pipe wall. For a turbulent jet discharging perpendicularly into a fully-developed turbulent pipe flow, the flow can be divided into three regions along the pipe: a flowestablishment region, a near-field region and a far-field region (Ger, 1974; Ger and Holley, 1976). Jet mixing dominates the transport mechanism in the first two regions, whereas the spread of the tracer in the far field is governed by the velocity and turbulent diffusion of the pipe flow. For Newtonian fluids, a number of experimental studies have been performed to study jet mixing in the first two of these regions for various jet-to-pipe velocity and diameter ratios, with either air or water as the working fluid and different measurement techniques, as summarized in Table 4.1. Three main mixing criteria have been utilized to characterize mixing: (a) standard deviation normalized by the cross-sectional mean value of concentration or temperature; (b) geometrically symmetric flow, with jet reaching the centre; and (c) pipe length required for the disappearance of colour from a base indicator or for temperature approaching equilibrium. Most past studies were conducted to determine the optimum relationship between jet-to-pipe velocity ratio and diameter ratio for different working fluids and mixing criteria. For the first criterion, the optimum jet-to-pipe velocity ratio minimizes the standard deviation of measured values of a tracer at a specified distance downstream of the injection point. The optimum mixing for the second criterion occurs when the jet penetrates to the centre of the pipe or the jet concentration profile becomes symmetric around the pipe axis at a specified distance at least two pipe diameters from the injection point (Forney, 1986). For the third criterion, optimum mixing was defined based on the shortest  54  distance required for neutralization of NaOH in the side stream, with HCl in the main stream (Cozewith and Busko, 1989). These correlations can be used to design the optimum flow conditions for specified diameter ratios of 90 side-tee mixers when the main stream and side stream are in the same phase (gas-gas or liquid-liquid) and both fluids are Newtonian. Although there are several studies on jet mixing in pipelines for Newtonian fluids, understanding jet mixing for non-Newtonian pulp suspensions is limited, since few studies have been conducted. Bobkowicz and Gauvin (1965, 1967) examined the distribution of hot water in nylon-fibre suspension flow in a vertical flow loop of diameter 50.8 mm. The tracer was injected at the pipe axis, and the radial temperature profile was measured by a temperature probe. Fibre diameters ranged from 20 to 52 m, their lengths from 0.52 to 1.2 mm, and fibre mass concentration from 0.5 to 6.0%. Radial dispersion, determined from the radial intensity of turbulence and eddy diffusion coefficient, increased substantially in the presence of fibres, even when drag reduction occurred. Andersson (1966) examined dye dispersion in a standard bleached spruce sulphite pulp flow in a rectangular channel, with turbulence generated by a grid at the inlet of the channel. The suspension concentration ranged from 0 to 5 kg/m3 and the velocity from 0.13 to 0.83 m/s. Dye dispersion in long-fibre pulp flow decreased with increasing concentration and was less than in comparable water flow. Luthi (1987) qualitatively investigated dispersion of dye injected at the centreline of a 6% pulp suspension flow with a mean velocity of 21.3 m/s. The dye persisted in the core 1.5 m  55  Table 4.1: Summary of previous studies on in-line jet mixing for Newtonian fluids (updated from Forney, 1986) Jet Mixing criterion  Jet fluid  Pipe fluid  Aq. NaCl  Pipe  Dj/D  Velocity ratio (R = Uj/Up)  Dj (cm)  Rej  D (cm)  Rep  Water  0.08 0.32  > 7,500  15.24  60,000  0.005 0.021  6 - 24  Aq. NaCl  Water  0.32  15.24  30,000 40,000  0.021  3 - 10.3  Aq. NaCl  Water  0.2  3.4  3,200  0.058  2.5 - 8.7  Air (25C)  Air (35C)  Air (25C)  Air (35C)  Air and CH4  Air  Air and CH4  Air  Air and CH4  Air  0.5 1.3 0.42 1.13 0.16 1.27 0.1 1.27 0.07 0.635  500 23,000 1,100 7,200 4,000 18,000  16,200 63,100 7,000 15,000 2,000 90,000 12,600 32,000 18,000 110,000  0.098 0.255 0.042 0.113 0.025 0.2 0.009 0.111 0.006 0.056  Pipe length required for indicator colour to disappear  Aq NaOH  Aq HCl  0.119 0.635  6,000 50,000  2,000 66,000  0.047 0.25  Length required for temperature to approach equilibrium  Hot water (50oC) Hot water (50oC)  Cold water (10oC) Cold water (10oC)  0.318, 0.635 0.318, 0.635  10,000 66,800 20,000 46,800  5,840 16,000 5,840 16,000  0.125, 0.25 0.125, 0.25  Conductivity/concentration standard deviation  Temperature standard deviation  Maximum concentration at the centre of the pipe  2,100 7,000 6,300 21,800 8,240 22,700 5,000  5.1 10 6.35 11.43 11.43 2.54 2.54 2.54  Measured variable Electrical conductivity Electrical conductivity Electrical conductivity  Reference Ger and Holley (1976) Fitzgerald and Holley (1981) Stephenson et al. (2007) Maruyama et al. (1981) Maruyama et al. (1983) Forney and Kwon (1979) Forney and Lee (1982) O'Leary and Forney (1985)  3-4  Temperature  3 - 10  Temperature  2-7  CH4 conc.  2.9 28.3  CH4 conc.  4.7 - 16  CH4 conc.  2 - 12  Colour  Cozewith and Busko (1989)  2.5 - 64  Temperature  Zughbi et al. (2003)  10 - 64  Temperature  Zughbi (2006)  56  downstream of the injection point, showing that plug flow occurred in the core of the pipe. Luettgen et al. (1991) measured cross-sectional conductivities of KCl solution injected into 0.5wt% hardwood pulp turbulent suspension flow via a side-injection port for jet-to-pipe velocity ratios of 0.75 - 12. The tracer concentration measured at several downstream positions declined more rapidly with increasing velocity ratio. They also found that KCl dispersed more rapidly in the pulp suspension than in water flow, and this behaviour occurred consistently for different velocity ratios. However, the results were limited since the experiments were conducted for only one fibre mass concentration (0.5%), and no mixing index was used to quantify the extent of mixing. Rewatkar et al. (2002) applied a temperature profiling technique to measure the mixing quality of pulp suspensions along the pipe axis downstream of a T-junction. They used the coefficient of variation (CoV) as a mixing index and found that the mixing quality downstream of tracer injection was significantly lower for the pulp suspensions than for water, due to suppression of downstream turbulence caused by interlocking fibres in the network. However, their experiments were conducted for only one mass concentration (1%) and one jet-to-pipe velocity ratio (4.2). More experimental data are clearly needed to provide better understanding of the mixing behaviour. Suspension rheology plays an important role in determining suspension behaviour. The complex rheology of pulp suspensions and its impact on the flow regime, contacts between fibres, floc formation, dispersion, etc. were summarized by Kerekes (2006). Due to their large aspect ratios (length-to-diameter ratio), fibres collide in rotation and transition in shear flow. Flocculation occurs when closely  57  spaced fibres collide and tangle with each other. Mason (1950) defined a condition where fibres were likely to collide in rotation with one fibre in the volume swept out by the length of a single fibre. The corresponding fibre volumetric concentration was called the “critical concentration,” given by  C v  1.5(d / L) 2  (4.1)  where L is the fibre length, and d is the fibre diameter. Based on this concept, Kerekes et al. (1985) defined the propensity for fibre flocculation in terms of a dimensionless number called the “crowding number”, defined as the number of fibres in a spherical volume of diameter equal to the fibre length, and given by  Nc   2 Cv ( L / d ) 2 3  (4.2)  For pulp fibre suspensions, it is more convenient to base the crowding number on fibre mass concentration, Cm (weight of fibres per weight of suspension, expressed in wt.%), and fibre coarseness,  (weight of fibre wall material for a specific fibre length, expressed in kg/m). Equation 4.2 can then be expressed as:  Nc   5C m L2    (4.3)  Mason’s critical concentration clearly corresponds to Nc = 1. When Nc < 1, each fibre is free on average to rotate without encountering other fibres, and flocculation generally does not occur. Corresponding fibre contacts were characterized as “chance collisions” by Soszynski (1987). Kerekes and Schell  58  (1992) showed that Nc can be directly related to the number of contacts per fibre in determining the suspension behaviour to form the networks. For Nc = 60, there are approximately 3 contacts per fibre, and fibres become restrained in rotation and locked into a network in a bent configuration (Meyer and Wahren, 1964). The 1 < Nc < 60 range was described by Soszynski (1987) as one of “forced collisions” between fibres. At higher crowding numbers, fibre mobility decreases significantly, and flocs in the suspension acquire mechanical strength through frictional forces among fibres. Martinez et al. (2001) divided the 1 < Nc < 60 range into two subregimes, separated by Nc = 16, defined as the “gel crowding number”. For Nc < 16, the fibres move freely, and the suspension behaves as essentially dilute. For 16 < Nc < 60, fibres interact with one another and flocculate, but fibre mobility persists. In recent work, Celzard et al.(2009) identified a “connectivity threshold” and “rigidity threshold” based on percolation and effective-medium theories, and showed that they correspond to Nc =16 and 60, respectively. The crowding number is a general indicator of the level of fibre contact in flowing systems, helpful in characterizing mixing quality in this study. ERT has been applied to study solid-suspension flow in horizontal pipes (e.g. Wang et al., 2003; Giguere et al., 2009). This technique has also been utilized extensively to evaluate mixing in various processes including pulp and paper systems, with details provided in sections 2.8. In this chapter, this technique has been applied to study mixing pulp fibre suspensions at tee mixers for various operating conditions to provide better understanding of jet mixing in pulp suspension flows.  59  4.2 Experimental Details The details of the test section and liquid injection system were described in section 3.1. The injection tube diameter was 3.81 mm (Dr = 0.05) for all experiments, except when otherwise identified. The tests were conducted for ranges of fibre mass concentrations (Cm = 0 – 3.0%), mainstream velocities (Up = 0.5 – 5.0 m/s) and side-stream velocities (Uj = 1.0 – 12.7 m/s), based on the procedure explained in section 3.2.1. The mixing quality was determined from measurements in cross-sectional planes along the pipeline based on the modified mixing index, described in section 3.3.1.  4.3  Results and Discussion  4.3.1 Jet mixing in Newtonian fluid (water) 4.3.1.1 Effect of velocities on mixing quality All tests were in the turbulent flow regime for both the main stream (Rep = 38,100 – 381,000) and the jet stream (Rej = 9,750 – 48,600). Figure 4.1 shows the effect of mainstream velocity on the mixing quality for virtually identical jet-to-pipe velocity ratio and the identical diameter ratio. Error bars in this and subsequent figures correspond to 90% confidence intervals. The mixing quality was almost independent of the mainstream velocity, with the small difference likely due to the retention time. The mixing quality was slightly better when flow rate was lower.  60  Tomographic images for the same condition are shown in Figure 4.2. From these images, the high-conductivity region in red and yellow represents the tracer or brine solution injected into the system, whereas the low-conductivity region in blue represents the main stream. The tracer was injected between planes 1 and 2. The mixing quality improved along the pipe, as can be seen from the disappearance of the red and yellow colours from planes 2 to 8. The jet penetration and the disappearance of the high-conductivity region were similar for each case, showing that the effect of mainstream velocity on mixing quality at the same velocity ratio was negligible.  100  Up(m/s) 10  velocity ratio (R)  1.5 2.0 3.0  P2  M' (%)  P3  P4  P5  4.50 4.41 4.23  P6  P7  1  P8  0.1 0  5  10  15  20  25  x/D Figure 4.1: Modified mixing index as a function of dimensionless distance downstream, x/D, for Newtonian fluid (water) at Up = 1.0 m/s, 2.0 m/s and 3.0 m/s for virtually identical jet-to-pipe velocity ratio and Dr = 0.05. Locations of planes P2 to P8 are shown in Figure 3.2.  61  (a)  mS/cm – milliSiemens per centimeter (b)  (c)  Figure 4.2: Tomographic images for Newtonian fluid (water) for Dr = 0.05 at (a) Up = 1.0 m/s, R = 4.50, (b) Up = 2.0 m/s, R = 4.41 and (c) Up = 3.0 m/s, R = 4.23. Locations of planes P1 to P8 are shown in Figure 3.2.  62  For tee mixers, the mixing quality of Newtonian fluids in turbulent flow strongly depends on jet penetration and thus on jet-to-pipe velocity ratio. For jet penetration in fully developed turbulent pipe flow, three mixing modes were defined by Pan and Meng (2001): wall-source, jet-mixing and jet-impaction. Figure 4.3 illustrates the effect of jet-to-pipe velocity ratio on the mixing quality at a jet-to-pipe diameter ratio, Dr, of 0.05. The mixing quality improved significantly as the jet-topipe velocity ratio increased. At low velocity ratio (R < 4), the jet momentum was insufficient for the jet stream to escape from the pipe wall, and it therefore became near-wall injection (i.e. wall-source mode). The mixing quality was then very poor. At higher velocity ratio (4 ≤ R ≤ 10), the jet momentum was sufficient for the jet to penetrate beyond the wall region and join the main flow. The modified mixing index was then significantly lower, i.e. the mixing quality was much better for the jetmixing mode. Further increases in velocity ratio (R > 10) allowed the jet to impinge on the far wall of the pipe. The jet then split and travelled around the pipe periphery, with mixing being efficient downstream due to the breakdown of large vortical structures from the jet impaction (Tosun, 1987). The jet penetrations in each scenario are shown in the tomographic images of Figure 4.4. For turbulent jets entering pipe flow from the side, jet mixing dominates the transport mechanism in the flow-establishment and near-field regions, whereas the velocity and turbulent diffusion of the pipe flow are dominant in the far field (Ger, 1974; Ger and Holley, 1976). The variation in concentration of tracer within a pipe cross-section decays exponentially with distance downstream in the far-field region, due solely to turbulent mixing in pipe flow, with negligible effect of jet mixing from  63  the first two regions. A plot of variance of tracer concentration versus distance (log:normal scale) is therefore first order only in the far-field region, >30 pipe diameters downstream. The plots in this study do not show first order behaviour, since the experimental results are for the jet-establishment and near-field regions, with the furthest measurement distance of only 22 pipe diameters.  100 R  M' (%)  10  P2  P3  P4  P5  P6  P7  1.36 2.16 3.18 4.38 6.39 6.93 8.28 9.92 12.5 15.5 17.8 25.3  P8  1  0.1 0  5  10  15  20  25  x/D Figure 4.3: Modified mixing index as a function of dimensionless distance downstream of injection for various jet-to-pipe velocity ratios, R, for water flow at Dr = 0.05.  The results from this work were also compared with those from previous work. Figure 4.5 shows the jet-to-pipe velocity ratios for the jet reaching the axis of the pipe at various jet-to-pipe diameter ratios for 90-tee mixers. The results are in reasonable agreement, despite the different experimental techniques adopted by different groups.  64  (a) Top  (b)  (c)  Figure 4.4: Tomographic images for (a) near-wall (wall-source) injection with R = 2.16, (b) jet-mixing with R = 6.39 and (c) jet-impaction with R = 15.5, Dr = 0.05 in water flow.  65  30 Forney and Kwon (1979) O'Leary and Forney (1985) Maruyama et al. (1983) Cozewith and Busko (1989) This study  R  10  1 0.03  0.6  0.1  Dr Figure 4.5: Jet-to-pipe velocity ratios, R, as a function of jet-to-pipe diameter ratio for the jet reaching the axis of the pipe for 90-tee mixers.  4.3.1.2 Effect of injection tube length on mixing quality Figure 4.6 shows the effect of injection tube length on mixing quality for the same  mainstream  and  side-stream  velocities.  The  entrance  length  was  approximately 24Dj for Uj = 2.0 m/s and Rej = 25,710. The longer injection tube length was sufficiently long (length: diameter ratio, Lj/Dj, of 36), to provide fully developed turbulent flow, whereas the flow was not fully developed for the short tube (Lj/Dj = 14). The modified mixing index for the fully developed flow was significantly lower than for the non-fully-developed one, suggesting better mixing when the side-entering jet was fully developed.  66  Figure 4.6: Comparison of long and short injection tubes on modified mixing index at Up = 1.0 m/s, Uj = 2.0 m/s, Dr = 0.167 for fibre-free water.  Tomographic images for the same condition are shown in Figure 4.7. The tracer was injected between planes 1 and 2. Mixing quality improved along the pipe, as seen from the disappearance of the red and yellow colours from plane 2 to plane 8. The fully developed side-stream flow penetrated further into the main pipe than the developing one since the velocity profile was flatter when the flow was not yet fully developed. Further penetration of the jet stream into the main stream provided better downstream mixing. Long injection tubes were therefore used for all subsequent tests.  67  (a)  (b)  Figure 4.7: Tomographic images for water flow at Up = 1.0 m/s, Uj = 2.0 m/s, Dr = 0.167: (a) short injection tube; (b) long injection tube.  4.3.1.3 Effect of temperature on mixing quality The temperature was varied from 15 to 25C, changing the viscosity of the water by 22%. However, the Reynolds number was high enough that the dependence of mixing quality on Reynolds number, and hence on liquid viscosity, was negligible.  68  4.3.2 Flow regime of pulp suspension flow in the pipe For Newtonian fluid flow through pipes, Reynolds numbers can be used to estimate the flow regime for specific test conditions. For pulp suspensions, however, characterizing flow regimes such as laminar, transition or turbulent is impossible since pulp suspension mostly flows as a plug, except near the walls where shear-dominated flow exists. Thus, it is difficult to ascribe a Reynolds number in a conventional sense, since the apparent viscosity of pulp suspension depends strongly on the shear rate, and shear is confined to a very small region near the pipe walls. A pulp suspension behaves as a solid at stresses less than the yield stress, whereas it exhibits shear-thinning for stresses exceeding the apparent yield stress. Further increasing the shear stress causes the suspension to behave as a Newtonian fluid (Derakhshandeh, 2011). Since the shear stress increases with radial distance from the centre of the pipe, a core region, in which the stress is less than the suspension yield stress, has solid-like properties and moves as a rigid plug, whereas the outer region, where the stress exceeds the yield stress, behaves like a liquid between the plug and the pipe wall. Once the shear stress exceeds the yield stress throughout the suspension, the plug disappears, and the flow becomes turbulent. For pulp suspensions in pipes, it is customary to identify the flow regime as plug, mixed or turbulent based on criteria suggested by Robertson and Mason (1957). Figure 4.8 shows a typical logarithmic head loss-velocity curve for lowconsistency pulp suspension flow through pipes. A plug flow regime exists in the lower velocity range, A to C. In the first part of this regime, A to B, friction increases  69  due to increased shear and increased frictional contact between pulp and wall. The suspension travels as a plug (referred to as “plug flow”). At point B, fibre migration away from wall commences and a clear water annulus forms between the plug and wall. This annulus increases in size with velocity, thereby decreasing the wall shear in a manner that offsets the increase in shear with increasing velocity alone. The water layer becomes turbulent at point C, which marks the onset of a mixed flow regime, in which turbulent fluid stresses on the plug surface exceed the pulp yield stress. Increasing velocity beyond C increases the size of the turbulent annulus by pulling additional fibres from the plug into the turbulent layer. At some velocity in the mixed regime, the friction loss of the suspension becomes less than for water, i.e. the flow becomes drag-reducing. Here fibres dampen the turbulence in the annulus, leading to lower friction loss in the suspension than would occur for water alone (Duffy and Lee, 1978). The mixed regime persists beyond D, with the plug decreasing in size, until the suspension flow becomes turbulent at point E. As one might expect, flow behaviour is strongly related to the rheological characteristics for pulp suspensions, in particular to pulp suspension yield stress. Rheological characteristics for pulp suspensions almost identical to those examined in the present work have been reported in previous work (Derakhshandeh, 2011). Figure 4.9 illustrates the relation between head loss and velocity for the softwood kraft pulp used in this work. For dilute pulp suspensions (Cm ≤ 1.0%), the flow can be either plug or turbulent, depending on the main stream velocity. At higher fibre mass concentrations (Cm ≥ 2.0%), however, the flow was essentially plug for all mainstream velocities investigated (0.5 - 5.0 m/s). For 0.5% pulp  70  suspension flow at Up ≤ 1.0 m/s, the imposed shear was less than the suspension yield stress in the core of the flow, leading to plug formation, with a thin dilute annulus in the mixed flow regime. Beyond this point, the plug disintegrated, and the flow became turbulent at Up ≥ 3.0 m/s. The head loss in the 0.5% pulp suspension became lower than in water between Up = 0.5 – 1.0 m/s; the intersection with the water curve reflects the onset of drag reduction for this fibre mass concentration. Drag reduction occurred at a higher velocity as the mass concentration increased, since more shear was required to disrupt the plug. For the highest mass concentration of 3.0%, however, the imposed shear was insufficient to disintegrate the plug, with friction loss being higher than in water for all mainstream velocities investigated (0.5 to 5.0 m/s).  Figure 4.8: Typical head loss-velocity curve for water and low-consistency pulp suspension flow through pipes (adapted from Robertson and Mason, 1957). Dashed line is the corresponding line for water.  71  Friction loss, H/L (m water/100 m pipe)  50 Cm = 3.0% Cm = 2.0%  10  Cm = 1.0%  water 1 Cm = 0.5% 0.3  1  10  Up (m/s)  Figure 4.9: Head loss-velocity curve for water and softwood pulp suspensions in this work.  Figure 4.10 illustrates the head loss curve for the hardwood kraft pulp. The flow again could be either plug or turbulent, depending on the mainstream velocity. At Cm = 1.0%, the flow was essentially plug in the mixed flow regime for Up ≤ 2.0 m/s, and became turbulent at Up ≥ 3.0 m/s. The onset of drag reduction was at Up  1.0 m/s. At the highest fibre mass concentration (Cm = 3.0%), the flow was essentially plug for all mainstream velocities investigated.  72  Friction loss, H/L (m water/100 m pipe)  50  Cm = 3.0% 10  Cm = 2.0%  Cm = 1.0% 1  water  Cm = 0.5% 0.3  1  10  Up (m/s)  Figure 4.10: Head loss-velocity curves for water and hardwood kraft pulp.  4.3.3 Effect of mainstream velocity on mixing quality in softwood pulp suspensions The results show significant differences in the jet mixing behaviour of nonNewtonian pulp suspensions and a Newtonian fluid (water). For pulp fibre suspensions, the mixing quality depends strongly on the flow regime, which varies with the mainstream velocity. Figure 4.11 shows the influence of the mainstream velocity on the mixing quality for pulp suspension at Cm = 0.5% at virtually identical jet-to-pipe velocity ratios and the identical diameter ratio. The effect of fibre networks on the mixing quality can be significant, even at a mass concentration as low as 0.5%. At a low velocity (Up = 1.0 m/s), turbulent shear was insufficient to  73  fluidize the suspension. The flow was therefore in the mixed flow regime with a plug in the core, suppressing downstream turbulence and hence decreasing the momentum transfer and mixing quality. At a higher velocity (Up = 2.0 m/s), the plug disintegrated, with the mixing quality improving significantly downstream. However, some plug likely persisted at this velocity, as shown by the suspension mixing quality being considerably lower than for water. Further increasing the velocity (Up ≥ 3.0 m/s) continued to disrupt the plug until it disappeared. The flow then reached the turbulent regime, and the mixing quality approached that for water.  100  M' (%)  Up(m/s)  R  1.0 3.63 2.0 3.36 3.0 2.96 4.0 3.16 4.0 (water) 3.17 (Rep =304,800)  10  Cm = 0.5%  P2 P3  P4 P5 Cm = 0% (water)  P6  P7  P8  1 0  5  10  15  20  25  x/D  Figure 4.11: Modified mixing index as a function of dimensionless distance downstream of injection for softwood pulp suspension at Cm = 0.5% and for water, Dr = 0.05 with various mainstream velocities and similar jet-to-pipe velocity ratios. The crowding number of 123 from equation 4.3 for Cm = 0.5% was higher than the critical value of 60, indicating > 3 contacts per fibre. At this condition, fibres  74  are locked into a bent configuration due to three-point contact. The resulting friction forces at fibre contacts give the network mechanical strength. For Up ≤ 1.0 m/s, this led to formation of a plug in the core, where the imposed shear was less than the suspension yield stress. For Up ≥ 3.0 m/s, the shear stress exceeded the yield stress throughout the suspension, providing relative motion among fibres and hence the turbulent fluidized regime in which floc breakdown due to shear stress was in dynamic equilibrium with floc formation (Robertson and Mason, 1957). This process was likely to suppress turbulent dispersion, causing mixing quality to be worse than in water. Tomographic images for the same conditions are shown in Figure 4.12. Clearly, jet penetration was similar for each case. The plug existed in the core of the pipe at a low velocity (Up = 1.0 m/s), with the high-conductivity region (tracer) persisting along the pipe as shown in Figure 4.12a. However, the high-conductivity region disappeared considerably faster downstream, i.e. the mixing quality significantly improved at higher mainstream velocities. For the turbulent suspension flow at Up = 4.0 m/s (Figure 4.12c), the distribution and disappearance of highconductivity region were similar to those for water (Figure 4.12d) at P2 and P5, but differed significantly downstream at P8. For the water flow, the tracer spread near the pipe wall, but it was concentrated further from the wall for the suspension flow, suggesting  that  floc  formation,  which  occurred  simultaneously  with  floc  disintegration by shear, likely suppressed turbulence in the core of the pipe.  75  (a)  Top  (b)  (c)  (d)  Figure 4.12: Tomographic images for softwood pulp suspension flow with Cm = 0.5%, Dr = 0.05 at the following: (a) Up = 1.0 m/s, R = 3.63; (b) Up = 2.0 m/s, R = 3.36; (c) Up = 4.0 m/s, R = 3.16; and (d) Newtonian fluid (water) at Up = 4.0 m/s, R = 3.17. Locations of planes P1, P2, P5 and P8 are shown in Figure 3.2.  76  At higher mass concentrations (Cm ≥ 2.0%), mixing was poor, even at a high mainstream velocity, since the imposed shear was not high enough to disrupt the plug and the suspension remained in plug flow. Figure 4.13 illustrates the effect of the mainstream velocity on the degree of mixing for a pulp suspension for Cm = 2.0% at the same diameter ratio and almost the same jet-to-pipe velocity ratios. The mixing quality improved only slightly as the mainstream velocity increased, and was not significantly better downstream due to a plug in the core, with considerably lower mixing quality than that for water. The strong fibre networks also affected jet penetration, although the jet-to-pipe velocity ratio was maintained nearly the same. As shown by the tomographic images in Figure 4.14, less jet penetration was observed at a low mainstream velocity due to the densely packed fibre networks.  100  Up (m/s)  M' (%)  Cm = 2.0%  1.0 3.77 2.0 3.36 3.0 3.04 4.0 3.22 4.0 3.17 (Rep = 304,800)  10  P2  Cm = 0% (water)  P3  P4  P5  P6  P7  P8  1 0  5  10  15  R  20  25  x/D  Figure 4.13: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension with Cm = 2.0% and for water, Dr = 0.05 for various mainstream velocities and similar jet-to-pipe velocity ratios.  77  (a)  Top  (b)  (c)  (d)  Figure 4.14: Tomographic images for softwood pulp suspension flow with Cm = 2.0%, Dr = 0.05 at the following: (a) Up = 1.0 m/s, R = 3.77; (b) Up = 2.0 m/s, R = 3.36; (c) Up = 4.0 m/s, R = 3.22; and (d) Newtonian fluid (water) at Up = 4.0 m/s, R = 3.17. Locations of planes P1, P2, P5 and P8 are shown in Figure 3.2.  78  4.3.4 Effect  of  jet  velocity  on  mixing  in  softwood  pulp  suspensions For a dilute mass concentration (Cm = 0.5%), the mixing behaviour in a pulp suspension was similar to that in water. The mixing quality depended strongly on the jet-to-pipe velocity ratio. Figure 4.15 shows the influence of the velocity ratio on the mixing quality for the wall-source and jet-mixing modes of mixing at Cm = 0.5% and Up = 2.0 m/s. The degree of mixing was similar for wall-source (R < 4) and better for jet-mixing (R = 6.26). The mixing quality improved considerably as the velocity ratio increased to reach jet-impaction, as shown in Figure 4.16. At this mainstream velocity (Up = 0.5 m/s), however, the mixing did not improve significantly downstream, likely due to decay of turbulence aided by reflocculation. At high velocity ratios (R = 17.6, 25.2), the jet momentum was sufficient to disrupt the plug when the jet was introduced to the main stream, as shown by improved mixing from P2 to P3, whereas mixing quality was not significantly better for lower velocity ratios (R = 8.11, 13.6). The modified mixing index then reached a plateau downstream, in contrast to that for water, for which the quality of mixing continued to improve. This was likely due to reflocculation, since the energy dissipation required to maintain turbulence from the side jet entry was not sustained.  79  M' (%)  100 Uj (m/s)  R  3.76 6.72 12.5 12.7  1.88 3.36 6.26 6.36  10  Cm = 0.5%  P2 P3 water  P4 P5  1 0  5  10  P6  P7  15  P8  20  25  x/D  Figure 4.15: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Cm = 0.5% and for water, Up = 2.0 m/s, Dr = 0.05 for various jet velocities in wall-source and jet-mixing modes. 100  Cm = 0.5% P2  M' (%)  10  P3  P4  P5  P6  P7  P8  Uj (m/s) 4.08 6.88 8.89 12.6 12.6  1  R 8.11 13.6 17.6 25.2 25.1  water 0.1 0  5  10  15  20  25  x/D  Figure 4.16: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Cm = 0.5% and for water, Up = 0.5 m/s, Dr = 0.05 for various jet velocities in jet-mixing and jet-impaction modes.  80  At higher Cm, a significant difference was found. Mixing depended on the penetration of the jet into the pipe. As shown in Figure 4.17, the mixing quality for a 2.0% consistency pulp suspension was very poor downstream at R = 13.5 when the jet attached to the far wall of the pipe. Tomographic images showing the jet penetration for this condition are illustrated in Figure 4.18b. The fibre networks were robust, and the energy dissipation from the jet impaction was not sufficient to disrupt the fibre networks, causing the jet to adhere to the far wall downstream. The mixing quality significantly improved at a higher velocity ratio (R =16.9) at which the jet momentum was sufficient to disintegrate the fibre networks and cause the jet to recirculate from the far wall to the core of the pipe, as shown in Figure 4.18c. However, the quality of mixing became worse at R = 24.9 when the jet attached to the opposite pipe wall (Figure 4.18d). This suggests that the mixing was optimal when the jet penetrated to the core of the pipe.  M' (%)  100  Uj (m/s)  R  4.88 6.88 8.64 12.7  9.62 13.5 16.9 24.9  P7  P8  10  P2  P3  P4  P5  P6  1 0  5  10  15  20  25  x/D  Figure 4.17: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Cm = 2.0%, Up = 0.5 m/s, Dr = 0.05 for various jet velocities in jet-mixing and jet-impaction mode.  81  (a)  Top  (b)  (c)  (d)  Figure 4.18: Tomographic images for softwood pulp suspension flow at Up = 0.50 m/s for Cm = 2.0%, Dr = 0.05 at the following: (a) R = 9.62; (b) R = 13.5; (c) R = 16.9; and (d) R = 24.9. Locations of planes P1, P2, P5 and P8 are shown in Figure 3.2.  82  4.3.5 Effect of velocities on mixing quality for hardwood pulp suspensions and optimum operating conditions for both pulp types The jet mixing behaviour of softwood and hardwood pulp suspensions was found to be very similar. The flow regime was important in determining the mixing quality for both pulp types. Figure 4.19 shows the effect of the mainstream velocity on the mixing quality for the hardwood pulp suspension at Cm = 1.0 and 3.0% for the same diameter ratio and almost the same jet-to-pipe velocity ratios. At Cm = 1.0%, the results were very similar to those for softwood pulp suspension at Cm = 0.5%. Increasing the mainstream velocity changed the flow regime from mixed flow (at Up ≤ 2.0 m/s) to the turbulent flow regime (at Up ≥ 3.0 m/s), resulting in plug disintegration and significantly better mixing, approaching that of the corresponding water flow, as shown in Figure 4.19a. At a higher mass concentration (Cm = 3.0%), however, (Figure 4.19b) the mainstream velocity had little effect on the mixing quality. The degree of mixing was poor, with only a slight increase in mixing quality observed with increasing mainstream velocity, since the fibre networks were robust and the suspension remained in plug flow for all mainstream velocities tested. Jet velocity is another important factor influencing mixing quality. The mixing behaviour was again similar for hardwood and softwood pulp suspensions. Figure 4.20 portrays the effect of jet velocity on mixing for hardwood pulp suspensions for Up = 0.5 m/s and Dr = 0.05. For dilute hardwood pulp suspensions (Cm ≤ 1.0%),  83  (a) 100  M' (%)  Up(m/s)  R  1.0 4.25 2.0 3.05 3.0 3.07 4.0 3.12 4.0 (water) 3.17 (Rep =304,800)  10  P2 P3  P4  P5  P6  P7  1 0  5  10  15  P8  20  25  x/D  (b)  M' (%)  100  Up (m/s)  R  2.0 3.0 4.0 4.0 (water)  3.06 3.05 3.11 3.17  10  Cm = 3.0%     water  1 0  5  10  15  20  25  x/D Figure 4.19: Modified mixing index as a function of dimensionless distance downstream at different fibre mass concentrations of hardwood pulp suspensions: (a) Cm = 1.0% and (b) Cm = 3.0% for Dr = 0.05, almost identical jet-to-pipe velocity ratios and various mainstream velocities, and comparison with modified mixing index for water under very similar experimental conditions.  84  the mixing quality depended strongly on the jet velocity. Mixing improved significantly with increasing jet velocity at Cm = 1.0%, as shown in Figure 4.20a. At Cm = 3.0% (Figure 4.20b), the mixing quality depended on the jet penetration into the pipe. The mixing quality was poor when the jet attached to the wall of the pipe (far wall at R = 12.3 and near wall at R = 24.1); the mixing was better when the jet penetrated to the core of the pipe (R = 8.19). The mixing was optimal at R =16.0 when the jet impinged on the far wall and circulated back to the core of the pipe. Tomographic images showing the jet penetration for each condition are illustrated in Figure 4.21. The data consistently showed this behaviour for Cm ≥ 2.0% (hardwood pulp) and Cm ≥ 1.0% (softwood pulp). However, the jet velocities or velocity ratios required for optimum mixing depended on the mass concentrations due to different fibre network strength, with higher velocity ratios needed for higher mass concentrations. The optimum values are summarized in Table 4.2. Decaying turbulence likely occurred at R = 16.0 as shown in Figures 4.20b and 4.21c. Energy dissipation from the jet impaction disrupted the fibre networks, and mixing quality improved for the first three planes (P2 – P4) downstream of the injection point as shown in Figure 4.20b. However, reflocculation probably occurred downstream, since the energy required to maintain turbulence was not sustained, causing mixing to worsen downstream. With the downstream reflocculation, a plug formed at the centre of the pipe, pushing the brine solution toward the pipe wall. This can be seen from the formation of the high-conductivity regions in red and yellow around the pipe, especially at the top of the P8 image in Figure 4.21c. This explanation also applies at R = 12.3 where the jet again disrupted the plug and  85  reached the far wall of the pipe, with reflocculation downstream. For R = 8.2, however, the tracer did not reach the opposite pipe wall since the jet fluid joined the main flow in the core of the pipe and travelled with the plug. (a) 100  Cm = 1.0% P2  M' (%)  10  P3  P4  P5  P6  P7  P8    1  Uj(m/s)  R  4.26 6.26 8.14 10.1 12.5 12.6 (water)  8.45 12.5 16.1 20.0 24.9 25.1  water 0.1 0  5  10  15  20  25  x/D  (b) 100  Uj(m/s)  M' (%)  4.13 6.26 8.26 12.5  R 8.19 12.3 16.0 24.1  10  1 0  5  10  15  20  25  x/D Figure 4.20: Modified mixing index as a function of dimensionless distance downstream at different fibre mass concentrations of hardwood pulp suspensions: (a) Cm = 1.0% and 0% and (b) Cm = 3.0% for Up = 0.5 m/s, Dr = 0.05 and various jet velocities.  86  (a)  Top  (b)  (c)  (d)  Figure 4.21: Tomographic images for hardwood pulp suspension flow at Up = 0.50 m/s for Cm = 3.0%, Dr = 0.05 at: (a) R = 8.2; (b) R = 12.3; (c) R = 16.0; and (d) R = 24.1. The locations of planes P1, P2, P5 and P8 are shown in Figure 3.2.  87  Table 4.2: Velocity ratios, R, for optimum mixing at Up = 0.5 m/s, Dr = 0.05 Pulp type Softwood  Hardwood a  Fibre mass concentration, Cm (%) 1.0 2.0 3.0 2.0 3.0  Optimum velocity ratio 10-14a 16.9 19.7 12-15a 16.0  Approximate; clearly between these two limits. In addition to mass concentration, the jet velocity (or velocity ratio) required  for the jet to reach the axis of the pipe is a function of the mainstream velocity since the fibre network strength depends on the mainstream velocity. At the same mass concentration, a higher velocity ratio is needed for lower mainstream velocity due to stronger fibre networks. This differs from the behaviour in water flow, as the velocity ratio, R, required for the jet to reach the centre of the pipe is almost independent of the mainstream velocity. The velocity ratios required for the jet to penetrate to the centre of the pipe for different pulp types, mass concentrations and mainstream velocities are summarized in Table 4.3. Due to a limitation of the liquid injection system, data could only be obtained for Up ≤ 2.0 m/s. At Up = 2.0 m/s, the required velocity ratio for the jet to penetrate to the axis of the pipe approached that for water, except for 3.0% softwood pulp suspension where the required velocity ratio was slightly higher than for water, suggesting that the effect of fibre network strength became less significant at this mainstream velocity. At higher velocities (Up > 2.0 m/s), therefore, the required velocity ratio also likely approaches that for water for the fibre mass concentration range investigated. Although the velocity ratios for optimum mixing in Table 4.2 are useful indicators, they are not necessarily desirable  88  from a practical point of view since impingement on the pipe wall creates significant stress there. At higher mass concentrations (Cm ≥ 2.0% for hardwood pulp suspension and Cm ≥ 1.0% for softwood pulp suspension), the mixing quality was poor when the jet attached to the pipe wall, and the design criteria can thus be based on the velocity ratio for the jet to reach the axis of the pipe, as shown in Table 4.3.  4.3.6 Effect of fibre mass concentration on mixing quality Figure 4.22 shows the effect of fibre mass concentration (consistency) on the mixing quality for Up = 4.0 m/s. The mixing quality improved considerably as Cm decreased. At this mainstream velocity, the pulp flow regime varied from plug to turbulent flow for different mass concentrations, leading to significant differences in mixing quality. At Cm ≥ 2.0, the flow was essentially plug, with densely packed fibre networks, and the degree of mixing was poor. At a lower concentration (Cm = 1.0%), disintegration of the plug was significant, and the mixing quality improved substantially. At Cm = 0.5%, the flow approached fully turbulent, and the degree of mixing was similar to that of water. The mass concentration did not, however, have such a strong effect on the mixing quality at the lower mainstream velocity of 1.0 m/s, as shown in Figure 4.23.  At this velocity, the degree of mixing improved  slightly as the mass concentration decreased since plug flow persisted with different fibre network strength. The mixing quality clearly differed significantly from that of water, even for Cm as low as 0.5%.  89  Table 4.3: Velocity ratios for jet reaching centre of pipe for Dr = 0.05 Fluid Water Softwood  Fibre mass concentration, Cm (%) 0.5  1.0  2.0  3.0  Hardwood  0.5  1.0  2.0  3.0  a  Mainstream velocity, Up (m/s) 0.5 1.0 2.0 0.5 1.0 2.0 0.5 1.0 2.0 0.5 1.0 2.0 0.5 1.0 2.0 0.5 1.0 2.0 0.5 1.0 2.0 0.5 1.0 2.0  Velocity ratio, R, for jet to reach axis of pipe 6.2 <8.0a 6.0 6.3 7.7 6.2 6.3 9.6 7.7 6.4 11.0 8.3 >6.3a <8.0a 6.2 6.2 <8.0a 6.2 6.2 8.2 6.3 6.2 9.0 8.1 6.2  Unable to obtain due to limitation of liquid injection system.  90  M' (%)  100  Cm = 3.0%  10  Cm = 2.0% Cm = 1.0%  P2 P3  P4  P5  Cm = 0.5%  P6  P7  1 0  5  10  15  water  P8  20  25  x/D  Figure 4.22: Modified mixing index as a function of dimensionless distance downstream at Up = 4.0 m/s, R ≈ 3.2 for various fibre mass concentrations.  100  M' (%)  Cm(%) 0 0.5 1.0 2.0 3.0  10  R 3.88 3.63 3.99 3.77 4.13  P2 P3  water  P4  1 0  5  P5  10  P6 15  P7 20  P8 25  x/D  Figure 4.23: Modified mixing index as a function of dimensionless distance downstream for Up = 1.0 m/s, virtually identical velocity ratios and various fibre mass concentrations.  91  4.3.7 Effect of fibre-turbulence interactions on mixing quality for hardwood pulp Figure 4.24 shows the effect of the mainstream velocity on mixing quality for a hardwood pulp suspension with Cm = 0.5% at virtually identical jet-to-pipe velocity ratios and identical diameter ratio. The suspension flow was clearly in the turbulent regime for all mainstream velocities examined, and the mixing quality in hardwood suspension flow was somewhat better than for water without fibres. The slight increase in mixing quality at a lower velocity (Up = 2.0 m/s) was likely due to the longer retention time.  M' (%)  100  10  Up(m/s)  R  2.0 3.0 4.0 4.0 (water)  3.06 3.09 3.08 3.17  P7  P8  P2 P3  P4  P5  1 0  5  10  P6 15  20  25  x/D Figure 4.24: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension at Cm = 0.5%, R = 3.1 compared with water for various mainstream velocities and almost identical jet-to-pipe velocity ratios.  92  Figure 4.25 compares the hardwood suspension at Cm = 0.5% and water flow at Up = 3.0 m/s for various jet velocities. Mixing quality improved with increasing velocity ratio, with consistently better mixing quality for the dilute hardwood suspension than for water in the turbulent flow regime.  100  M' (%)  Uj (m/s) 4.07 4.13 9.60 9.26 12.7 12.3  10  P2  R 1.36 (W) 1.38 (HW) 3.19 (W) 3.09 (HW) 4.23 (W) 4.11 (HW)  P3 P4 1 0  5  P5  10  P6 15  P7 20  P8 25  x/D Figure 4.25: Modified mixing index as a function of dimensionless distance downstream for water (W) and hardwood pulp suspension (HW) flow at Cm = 0.5%, Up = 3.0 m/s with various jet velocities.  However, mixing quality was worse for the softwood suspension at the same fibre mass concentration in the turbulent flow regime, as shown in Figure 4.26. This might be due to the longer fibres creating stronger fibre networks for the softwood than for the hardwood, thereby suppressing turbulent fluctuations and hence mixing. Another factor could be increased resistance to bending of the larger-diameter softwood fibres when subjected to flow-induced stresses, affecting the extent to  93  which the two types of fibres dampen turbulence. The differences in the mixing quality are relatively small, but significant, since error bars for each curve seldom overlap.  100  M' (%)  Softwood fibers Hardwood fibers Water (no fibers) 10  P2 P3  P4  P5  1 0  5  10  P6 15  P7 20  P8 25  x/D Figure 4.26: Modified mixing index as a function of dimensionless distance downstream for water, softwood and hardwood pulp suspensions at Cm=0.5%, Up = 4.0 m/s, R = 3.1 and Dr = 0.05.  The improved mixing for the short-fibre hardwood pulp suspension coincides with results reported by Luettgen et al. (1991), who also found that turbulent dispersion was greater for a dilute hardwood suspension (at Cm = 0.5%) than for fibre-free water. The results also agree with findings summarized by Bobkowicz and Gauvin (1965, 1967), who reported that the radial intensity of turbulence in water flow containing short nylon fibres (of length from 0.52 to 1.21 mm) was higher than for water without fibres, even though the suspension flow was in the drag reduction  94  regime. It is generally believed that drag reduction is due to the suppression of turbulent eddies by fibres or other particles. Several studies (Lee et al., 1974; Vaseleski and Metzner, 1974; Lee and Duffy, 1976a, 1976b; Sharma et al., 1979) have suggested that drag reduction in fibre suspensions is especially due to dampening of radial velocity fluctuations in the turbulent core, in contrast to drag reduction in polymer solutions, in which drag reduction is due to dampening near the wall (Virk, 1975; McComb and Rabie, 1982). This implies less turbulent dispersion and reduced lateral mixing, which appears to contradict the experimental results of improved mixing for the hardwood fibres. However, fibre aspect ratios in those previous studies were relatively high (>70), close to that for the softwood fibres (92) and significantly higher than for the hardwood fibres (36) in this study. A softwood suspension likely forms fibre networks, dampening turbulent fluctuation in the core, whereas a short-fibre hardwood suspension probably behaves differently. The crowding number of 123, calculated from equation 4.3 for the 0.5% softwood suspension, significantly exceeded the critical value of 60, indicating more than 3 contacts per fibre. Consequently, fibres entangled, flexed and formed networks from frictional forces between them. In the turbulent fibre suspension flow, these networks were in a dynamic equilibrium between simultaneous floc formation and breakdown of flocs by shear stress (Robertson and Mason, 1957), and likely suppressed turbulent dispersion in the core of the suspension flow, leading to less momentum transport than in water flow and hence to drag reduction. For the hardwood, the crowding number of 13.8 for Cm = 0.5% was lower than the gel crowding number of 16. Accordingly, fibres were free to move relative to one  95  another. Individual fibres were then free to migrate throughout the main flow, and the mechanism of drag reduction may have differed from that for softwood fibres. The changes in turbulence structure causing drag reduction in hardwood fibres probably occur only in the wall layer, as for drag-reducing polymeric fluids. Figure 4.27 compares tomographic images for water (Cm = 0) with Cm = 0.5% softwood and hardwood suspensions for the same flow conditions in the turbulent flow regime. Similar jet penetration was observed at P2 for each case. The distribution and disappearance of high conductivity regions in the hardwood suspension were similar to those in water, but differed from those in the softwood suspension, as shown from the image contours downstream, especially at P6 and P8. The drag reduction mechanism for the hardwood fibres therefore probably differed from that for the softwood fibres. Due to high number of fibre contacts in the softwood suspension suggested by the crowding number, floc formation likely suppressed turbulence in the core of the pipe, consistent with the high-conductivity region being concentrated further from the pipe wall at P6 and P8. In the hardwood fibres with lower crowding, however, the tracer spread near the pipe wall, indicating negligible floc formation occurring in the core, and turbulence was probably dampened in the near-wall region. Accordingly, drag reduction in hardwood suspensions might not be a determining factor for the overall mixing quality, since it is predominantly related to near-wall behaviour, whereas mixing in our study was measured over the entire pipe cross-section. This implies that short fibres might alter the turbulence structure  96  (a)  Top  (b)  (c)  Figure 4.27: Tomographic images for Cm=0.5%, Up = 4.0 m/s, R = 3.1 and Dr = 0.05 with: (a) softwood; (b) hardwood; and (c) fibre-free water. The locations of planes P2, P4, P6 and P8 are shown in Figure 3.2. in the bulk, leading to better mixing in hardwood suspensions than in fibre-free water. The crowding number for the hardwood fibres suggested that the suspensions behaved as essentially dilute, with individual fibres moving freely of one another throughout the flow, possibly carrying turbulent eddies with them.  97  Fibres may also have collided with each other, generating local turbulent eddies. These factors could promote turbulent dispersion, resulting in improved mixing in a dilute hardwood suspension. McComb and Chan (1985) reported that the longitudinal component of turbulent velocity fluctuations in fibre suspension flow decreased, whereas the tangential component increased, relative to water flow. An increase in tangential fluctuations could promote turbulent dispersion and mixing quality. Luettgen et al. (1991) proposed that the motion of flocs in the suspension flow provided a mechanism for increasing longitudinal and tangential velocity fluctuations, while suppressing radial fluctuations, and proposed that this mechanism could promote turbulent dispersion. Figure 4.28 plots the modified mixing index for various fibre mass concentrations at Up = 4.0 m/s. The results show the effects of flow regime and fibre-turbulence interactions on the mixing quality. Plug flow was approached at Cm ≥ 2.0%, leading to poor mixing downstream. At lower mass concentrations (Cm ≤ 1.0%), the flow was in the turbulent regime, and the mixing quality was much better. Better mixing quality than for water occurred at Cm = 0.5%, but not at Cm = 1.0%. The crowding number was 13.8 for Cm = 0.5%, whereas it was 27.6 for Cm = 1.0%. The latter was higher than the gel crowding number of 16, causing fibres to interact and begin to flocculate. It appears that the fibre mass concentration needs to be sufficiently high for fibre-turbulence interactions to improve dispersion, but not so high as to engender fibre networks strong enough to reduce turbulent fluctuations.  98  100  M' (%)  Cm(%) 0 0.5 1.0 2.0 3.0  10  P2 P3  P4  P5  P6  1 0  5  10  15  P7 20  P8 25  x/D Figure 4.28: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension at Up = 4.0 m/s, R = 3.1 with various fibre mass concentrations.  4.3.8 Effect of fibre type on mixing quality Figure 4.29 shows the effect of fibre type on mixing at Up = 3.0 m/s. The mixing quality for the short-fibre hardwood pulp suspension was substantially better than for the softwood suspension, and was very similar to that for water, as illustrated in Figure 4.29a. At this consistency and mainstream velocity, the turbulent shear was sufficient to disrupt the fibre networks for hardwood suspension, causing the flow to become turbulent, leading to enhanced mixing. For the same conditions, the fibre networks for softwood suspension were robust, flowing as a plug, resulting in poor mixing. The fibre properties did not, however, have such a strong effect on mixing quality at the higher mass concentration of  99  3.0%, as indicated in Figure 4.29b. In this case, mixing was similar for both fibre types, and poor relative to water, since plug flow occurred in both cases. Supplementary data relevant to this chapter are provided in Appendix A.  4.4 Conclusions In-line jet mixing was investigated based on non-intrusive electrical resistance tomography for a range of operating conditions for water, softwood and hardwood kraft pulp suspensions. For water in the turbulent flow regime, the quality of mixing improved substantially when the jet-to-pipe velocity ratio increased, changing the mixing mode from wall-source to jet-mixing and jet-impaction. At the same jet-to-pipe velocity ratio, jet penetration and mixing quality were almost independent of the mainstream velocity for a given jet-to-pipe diameter ratio. The mixing quality for a long injection tube providing fully developed jet flow was significantly better than for a short one due to the flatter velocity profile of developing flow causing less jet penetration. Temperature variation between 15 and 25°C had insignificant effect on the mixing quality. For pulp fibre suspensions, the flow regime, which varies with the mainstream velocity, played an important role in determining the degree of mixing. For dilute suspensions (Cm = 0.5% for softwood pulp suspension and Cm ≤ 1.0% for hardwood pulp suspension), mixing quality improved profoundly with increasing mainstream velocity. At a low velocity (0.5 m/s), the flow was in the mixed flow regime. A rigid plug formed in the core of the flow, where the shear stress was less  100  (a) 100  M' (%)  Softwood fibers Hardwood fibers Water (no fibers) 10  P2 P3  P4  P5  P6  P7  P8  1 0  5  10  15  20  25  x/D (b) 100  M' (%)  Softwood fibers Hardwood fibers Water (no fibers) 10  1 0  5  10  15  20  25  x/D Figure 4.29: Modified mixing index as a function of dimensionless distance downstream for water, softwood and hardwood pulp suspensions at (a) Cm = 1.0% and (b) Cm = 3.0%, for Up = 3.0 m/s and R = 3.1.  101  than the suspension yield stress, and a thin dilute annulus occurred between the plug and the wall, where the shear exceeded the yield stress. The strong fibre networks in the core dampened the turbulence, leading to poor mixing downstream. At higher velocity, the plug was disrupted, leading to much better mixing downstream. The quality of mixing approached that in water when the flow became turbulent at Up ≥ 3.0 m/s. Slightly lower mixing quality in turbulent suspension flow than in water was likely due to floc formation, in dynamic equilibrium with breakdown of flocs by turbulent shear, suppressing turbulence. At higher concentrations (Cm ≥ 2.0% for softwood pulp and Cm = 3.0% for hardwood pulp), mixing was poor, even at high mainstream velocities, since the imposed shear was insufficient to disrupt the plug, so that additional shear would be required to improve mixing. For plug flow in the mixed flow regime with a loose fibre network, the jet could disrupt the plug. Mixing improved with increasing jet velocity, but was significantly lower than for water for Cm = 0.5% with softwood pulp, and for Cm ≤ 1.0% with hardwood fibres. However, at higher fibre mass concentrations, the mixing quality depended on jet penetration. Jet attachment to the far wall of the pipe led to poor mixing. Mixing quality improved when the jet penetrated to the axis of the pipe, and improved further when the jet impinged on the opposite wall and then recirculated to the core of the pipe. Higher jet velocities were required for the jet to penetrate to the centre of the pipe at higher mass concentrations and lower mainstream velocities due to denser fibre networks.  102  In the turbulent flow regime, mixing for a dilute hardwood suspension at Cm = 0.5% was somewhat better than for water, whereas it was worse for a softwood suspension at the same concentration. Drag reduction may not be directly linked to the overall degree of mixing in hardwood suspensions since it predominantly reflects behaviour in the wall layer. Shorter smaller-diameter fibres likely modify the turbulence structure in the bulk, resulting in better mixing over the entire pipe crosssection, whereas high mass concentration and long large fibres create fibre networks that reduce turbulent fluctuations and mixing quality.  103  5. Mixing Liquid into Pulp Fibre Suspensions in the Presence of an In-Line Mechanical Mixer6  5.1 Introduction In-line mechanical mixers are primarily used in continuous processes. The mixers provide low holdup, valuable for rapid mixing to avoid poor product quality or formation of undesired by-products (Myers et al., 1999). For these devices, the residence time is low, and the agitators rotate with a high speed. One of the most popular types of in-line mechanical mixers is a rotor-stator mixer, consisting of a rotor, which spins at a high speed inside a fixed stator, providing short residence time and high energy dissipation rate. Design and characteristics of rotor-stator mixers were summarized by Atiemo-Obeng and Calabrese (2004). Early studies on in-line rotor-stator mixers were carried out to evaluate micromixing for fast azocoupling reactions (Bourne and Garcia-Rosas, 1986; Bourne and Studer, 1992) and to determine power characteristics for foam generation (Kroezen et al., 1988; Hanselmann and Windhab, 1999). Recently, Baldyga et al. (2007) investigated the break-up of nano-particle clusters in nano-suspensions by an in-line Silverson rotorstator mixer based on CFD. The results indicated strong dependence of particle 6  A version of this chapter has been accepted for publication: Yenjaichon, W., Grace, J.R., Lim, C.J.,  Bennington, C.P.J., 2012. Pilot-scale examination of mixing liquid into pulp fiber suspensions in the presence of an in-line mechanical mixer, Ind. Eng. Chem. Res.  104  breakage on rotor speed and suspension flow rate. Increasing the flow rate (less residence time) decreased the extent of particle break-up, especially at higher rotor speeds. The mean particle size did not decrease significantly with increasing rotor speed at higher flow rate, implying less influence of rotor speed at shorter residence times. Baldyga et al. (2008) showed that the model predictions agreed well with the experimental results of particle breakage by this in-line mixer. Cooke et al. (2012) developed an expression for the power consumption of a Silverson high-shear inline rotor-stator mixer to describe the impact of rotor speed and flow rate for Newtonian and shear-thinning non-Newtonian fluids. Hall et al. (2011a) investigated droplet break-up by this in-line mixer. The emulsion drop size was found to decrease significantly with increasing rotor speed, but was almost independent of the flow rate. Hall et al. (2011b) showed that increasing the flow rate increased the total energy dissipation rate, with a minimal change in drop size, whereas the drop size decreased significantly with increasing rotor energy dissipation rate, with residence times maintained similar. In bleaching processes, in-line mechanical mixers are common for mediumconsistency applications. They are usually installed in conjunction with jet mixers such as peg mixers and high-shear mixers, with details provided in section 2.2. The main objective of this chapter is to investigate the influence of mainstream velocity, jet velocity, impeller speed, fibre mass concentration and fibre properties on mixing of pulp suspensions with an in-line mechanical mixer, to provide better understanding and useful criteria for industrial applications.  105  5.2 Experimental Details A schematic of the in-line mechanical mixer is shown in Figure 3.5. The injection tube diameter was 3.81 mm (Dr = 0.05) for all the experiments covered in this chapter. The experiments were conducted using the procedure described in section 3.2.1 with mixing quantified by the modified mixing index, defined in section 3.3.1. The suspension concentrations varied from 0 to 3.0%, mainstream velocities from 0.5 to 3.0 m/s, jet velocities from 3.8 to 12.6 m/s and impeller speeds from 0 to 800 rpm. The mechanical mixing system was described in section 3.1. The injection tube is sufficiently long to provide fully-developed flow, with a length:diameter ratio of 47, compared with an entry length of approximately 27Dj for the highest sidestream velocity investigated.  5.3 Results and Discussion 5.3.1 Effect of velocities and impeller rotational speed on mixing quality in water All tests were in the turbulent flow regime for both the main stream (Rep = 38,100 – 229,000) and jet stream (Rej = 9,750 – 48,600). Figure 5.1 shows the effect of mainstream velocity on the mixing quality with the perpendicular static impeller and rotating impeller at N = 400 rpm and Dr = 0.05. At this rotational speed, the mixing behaviour for the rotating impeller was similar to that for the static impeller. At the same jet-to-pipe velocity ratio, the mixing quality improved  106  downstream and was almost independent of the mainstream velocity, indicating negligible effect of residence time on mixing. The data consistently showed the independence of the mainstream velocity at constant velocity ratios for both rotating and perpendicular static impellers, similar to findings for a tee mixer alone.  (a) 100  10  M' (%)  P2 P3  P4  1  P5  P6  Up(m/s)  R  1.0 2.0 3.0 1.0 2.0  4.13 4.13 4.17 6.10 6.13  P7  P8  Impeller 0.1 0  5  10  15  20  25  x/D  (b) 100  Up(m/s) R 1.0 2.0 3.0 1.0 2.0  M' (%)  10  4.03 4.13 4.12 6.11 6.20  1 Impeller 0.1  0  5  10  15  20  25  x/D  Figure 5.1: Modified mixing index as a function of dimensionless distance downstream for Newtonian fluid (water) at Up = 1.0, 2.0 and 3.0 m/s, Dr = 0.05 with: (a) perpendicular static impeller (b) impeller rotating at N = 400 rpm. Locations of planes P2 to P8 are shown in Figure 3.5.  107  Figure 5.2 illustrates the influence of mainstream velocity on the quality of mixing at higher rotational speeds. Improved mixing was observed with decreasing mainstream velocity, with considerably better mixing quality at Up = 1.0 m/s. The difference was greater at higher rotational speed since the influence of residence time in the high-shear zone around the rotating impeller on mixing was more significant at higher rotational speed. At a lower impeller speed (N = 400 rpm), however, the energy dissipation rate was lower, and the residence time in the highshear zone had insignificant influence on mixing, as shown in Figure 5.1b. The effect of impeller speed on the mixing quality for water at Up = 1.0 m/s is presented in Figure 5.3. Two configurations of the impeller when static ─parallel and perpendicular to the flow─ were also examined. The impeller speed varies from 400 to 800 rpm, equivalent to the tip velocity of 1.3 to 2.7 m/s. The mixing quality clearly improved when the impeller was added. The stationary impeller parallel to the flow (with a schematic shown as an inset in Figure 5.3) provided better mixing than the tee alone, whereas the mixing quality of water flow with the perpendicular orientation was significantly better than for the parallel one, and the mixing quality was similar to that for the mechanical mixer at N = 400 rpm. For this velocity, mixing improved profoundly as the rotational speed increased. The impeller speed and configuration had less influence on mixing at higher mainstream velocities as shown in Figure 5.4. At higher velocities, the fluid passing through the high-shear zone around the rotating impeller likely spent less time there. An increase in impeller speed therefore did not improve mixing significantly, with mixing quality similar to that of the perpendicular static mixer. At a lower  108  (a) 100  Up(m/s) R 1.0 2.0 3.0  10  M' (%)  P2  P3  P4  1  P5  P6  4.26 4.11 4.21  P7  P8  Impeller 0.1  0  5  10  15  20  25  x/D  (b) 100  Up(m/s) R 1.0 2.0 3.0  M' (%)  10  4.25 4.13 4.23  1  Impeller  0.1 0  5  10  15  20  25  x/D  Figure 5.2: Modified mixing index as a function of dimensionless distance downstream for water at Up = 1.0, 2.0 and 3.0 m/s, Dr = 0.05, virtually identical velocity ratios with: (a) N = 600 rpm (b) N = 800 rpm.  109  100  Parallel (//) Perpendicular ( )  M' (%)  10  flow     N (rpm) NA 0 (//) 0( ) 400 602 803  flow  R 3.88 4.13 4.13 4.03 4.26 4.25  P2  1  P3 Impeller  P4  0.1 0  5  P5  10  P6 15  P7 20  P8 25  x/D  Figure 5.3: Modified mixing index as a function of dimensionless distance downstream of injection for water at Up = 1.0 m/s, Dr = 0.05, almost constant velocity ratios and various rotation speeds.  mainstream velocity, however, the fluid stayed longer in the high-shear zone, with energy dissipation increasing with increasing impeller rotational speed, leading to significantly improved mixing. The results were similar to those from previous studies on an in-line rotor-stator mixer (Baldyga et al., 2007, 2008), where for a high-rotor speed, particle break-up occurred mainly around the inner rotor at a lowflow rate (long residence time), and around the outer rotor as well at a high-flow rate. For a low-rotor speed, however, particle breakage occurred in the region where the outer rotor operated for both low and high flow rates. This implies that the residence time is more significant at higher rotor speed. The results also showed that the mean particle size did not decrease significantly with increasing rotor speed  110  at higher flow rate, consistent with the present work, indicating less influence of the impeller speed on mixing at higher mainstream velocities.  (a) 100 N (rpm) R NA 0 (//) 0( ) 200 403 601 803  M' (%)  10  P2  1  P3  P4  P5  Impeller 0.1  0  5  10  P6  15  P7  4.41 4.13 4.13 4.13 4.11 4.13 4.13  P8 20  25  x/D  (b) 100  M' (%)  10  N (rpm)  R  NA 0 (//) 0( ) 404 602 805  4.23 4.12 4.17 4.12 4.21 4.23  1  Impeller 0.1 0  5  10  15  20  25  x/D  Figure 5.4: Modified mixing index as a function of dimensionless distance downstream of injection for water, Dr = 0.05, almost constant velocity ratios and various rotation speeds at: (a) Up = 2.0 m/s (b) Up = 3.0 m/s.  111  Figure 5.5 shows the influence of jet-to-pipe velocity ratio at N = 400 rpm. At this impeller speed, the residence time, and hence the mainstream velocity, had less effect on mixing, and the mixing quality depended strongly on the jet-to-pipe velocity ratio. The mixing quality improved significantly with increasing velocity ratio as the mixing mode changed from wall-source (R < 4) to jet-mixing (4 ≤ R ≤ 10) and jet-impaction (R > 10). The results for the in-line mechanical mixer were similar to those for a tee mixer described in section 4.3.1.1. However, a significant difference was observed at R = 12.2 when the jet stream reached the far wall of the pipe ahead of the impeller. The mixing quality was worse than for the jet penetrating to the centre of the pipe, behaviour not observed for the tee mixer alone. The mixing was more efficient when the jet penetrated to the core of the pipe, likely due to higher energy dissipation when the jet impinged on the rotating impeller. At very high velocity ratios (R > 16), the energy dissipation from jet impingement on the pipe wall and rotating impeller was profound, resulting in efficient downstream mixing. Tomographic images showing jet penetration and downstream mixing appear in Figure 5.6. The tracer was injected and reached the impeller between planes 1 and 2. The disappearance of the red and yellow colours from plane 2 to plane 8 showed improved mixing downstream. The jet penetrated to the core of the pipe at R = 6.15, reached the far wall of the pipe ahead of the impeller at R = 12.2, and impinged on the far wall and recirculated around the pipe periphery ahead of the impeller at R = 24.6.  112  100 R 2.06 3.00 4.10 6.15 8.21 12.2 16.5 24.6  M' (%)  10  1  P2 P3  0.1  Impeller  P4  P5 P6  0  5  10  15  P7 20  P8 25  x/D  Figure 5.5: Modified mixing index as a function of dimensionless distance downstream of injection for various jet-to-pipe velocity ratios, R, with water at Dr = .05, N = 400 rpm.  5.3.2 Effect of mainstream velocity on mixing in pulp suspensions Figure 5.7 plots the modified mixing index for various mainstream velocities for a softwood pulp suspension at Cm = 0.5%, N = 400 rpm, almost identical jet-topipe velocity ratios and identical diameter ratio. Unlike mixing in water at similar velocities and rotational speed, pulp suspension mixing depends strongly on the mainstream velocity because of flow regime differences. The mixing improved significantly when the flow regime changed, and approached that for water when the flow was turbulent for Up ≥ 2.0 m/s. The flow regimes for the softwood and  113  (a)  (b)  (c)  Figure 5.6: Tomographic images for water, Dr = 0.05, N = 400 rpm at: (a) R = 6.15 (b) R = 12.2 (c) R = 24.6. Locations of planes P1 to P8 are shown in Figure 3.5.  114  hardwood pulp suspensions in this study were discussed in section 4.3.2. Without an impeller, the suspension flowed as a plug at Up = 1.0 m/s, and was turbulent for Up = 3.0 m/s. At Up = 2.0 m/s, the plug was disrupted, but the flow was not yet turbulent without an impeller. The addition of the impeller was sufficient to introduce turbulence, with mixing quality approaching that for water. Without the impeller, the mainstream velocity required for the flow to become fully turbulent and for the mixing quality to approach that in water was as high as 4.0 m/s. In the high-shear zone (P2 in Figure 3.5) immediately downstream of the impeller, the flow was turbulent, with mixing quality in suspension being similar to that in water and independent of the mainstream velocity. Reflocculation likely occurred rapidly for Up = 1.0 m/s, suppressing downstream turbulence, as shown by considerably worse downstream mixing compared to that for Up ≥ 2.0 m/s. The crowding number, Nc, of 123 from equation 4.3 for the 0.5% softwood pulp suspension was higher than 60, indicating >3 contacts per fibre and flocculation when the imposed shear was less than the suspension yield stress. When the softwood suspension flow in an empty pipe was tested at Up = 3.0 m/s, the shear stress exceeded the suspension yield stress, providing relative motion among fibres and turbulent flow. At Up = 1.0 m/s, however, the flow was in the mixed flow regime, with a rigid plug in the core. The rotating impeller disrupted the plug and provided turbulence in the high-shear (P2) zone, with reflocculation likely occurring downstream since the energy dissipation was not sustained  115  100  Up(m/s) 1.0 2.0 3.0 3.0 (w)  M' (%)  10  P2  1  P3  P4  P5  Impeller  P6  P7  R 3.99 4.13 4.13 4.12  P8  0.1 0  5  10  15  20  25  x/D  Figure 5.7: Modified mixing index as a function of dimensionless distance downstream of injection for softwood pulp suspension with Cm = 0.5% and water (w), Dr = 0.05, N = 400 rpm with various mainstream velocities and almost identical jet-to-pipe velocity ratios.  Figure 5.8 shows the influence of the mainstream velocity on the degree of mixing of a softwood pulp suspension at Cm = 2.0% for virtually identical impeller speeds and jet-to-pipe velocity ratios and identical diameter ratio. At this concentration, the flow was essentially plug before passing through the impeller for all mainstream velocities investigated. The rotating impeller disrupted the plug, and mixing quality improved substantially with increasing mainstream velocity compared to the tee mixer alone, where mixing quality was poor and improved only slightly as the mainstream velocity increased (see Figure 4.13). Enhanced mixing with increasing mainstream velocity at this concentration occurred mainly in the high-  116  shear (P2) zone, whereas mixing improved only slightly downstream for Up ≥ 2.0 m/s compared to that for Cm = 0.5%, likely due to re-establishment of a plug downstream of the impeller. At Up = 3.0 m/s, the quality of mixing in the pulp suspension was considerably better with the impeller than without it at the same mainstream velocity. The mixing was slightly worse than for water at P2 and P3, but significantly worse downstream, suggesting that reflocculation likely occurred rapidly downstream of the impeller, with robust fibre networks impeding mixing.  100  Cm = 2.0% 10  M' (%)  Up (m/s)  P2  1  P3  Cm = 0% (water)  P4  P5  Impeller  P6  P7  R  3.0 4.23 1.0 4.09 2.0 4.13 3.0 4.13 3.0 (w) 4.12  N (rpm) 410 420 433 404  P8  0.1 0  5  10  15  20  25  x/D  Figure 5.8: Modified mixing index as a function of dimensionless distance downstream with and without impeller for softwood pulp suspension for Cm = 2.0% and water (w), Dr = 0.05, various mainstream velocities and almost identical jet-topipe velocity ratios and rotational speeds.  117  5.3.3 Effect of jet velocity on mixing quality in pulp suspensions Figure 5.9 portrays the effect of the jet-to-pipe velocity ratio on the mixing quality for a softwood suspension at Cm = 0.5%, Up = 1.0 m/s and N = 400 rpm. The mixing was significantly better when the mixing mode changed from wall-source (R = 3.4) to jet-mixing (R = 6.3). It was worse, however, at higher velocity ratio when the jet reached the far wall of the pipe, whereas this behaviour was not observed at the same mass concentration for a tee mixer alone. The difference was likely due to faster energy dissipation when the jet penetrated to the core of the pipe and impinged on the rotating impeller at R = 6.3, as discussed above for water flow.  100  R  N (rpm)  3.99 6.26 8.24 12.4  M' (%)  10  402 403 403 402  P2  1  P3  P4  P5  Impeller  P6  P7  P8  0.1 0  5  10  15  20  25  x/D  Figure 5.9: Modified mixing index as a function of dimensionless distance downstream of injection for softwood pulp suspension with Cm = 0.5%, Up = 1.0 m/s, Dr = 0.05 and various jet velocities.  118  Figure 5.10 shows the mixing quality for a softwood suspension with Cm = 0.5%, Up = 0.5 m/s and higher velocity ratios. Mixing improved significantly with increasing jet velocity, likely because jet impaction and the rotating impeller disintegrated the plug in the dilute pulp suspension. The energy supplied was not, however, enough to provide turbulence, even in the high-shear (P2) zone, with reflocculation occurring downstream, as shown by considerably worse mixing than for water beyond P3.  100  Cm = 0.5%  M' (%)  10  1  R  N (rpm)  12.4 16.6 24.7 24.6  403 403 404 402  P2 P3  0.1  P4  Cm = 0% (water)  P5 P6  Impeller  P7  P8  0.01 0  5  10  15  20  25  x/D  Figure 5.10: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension with Cm = 0.5% and water (w), Up = 0.5 m/s, Dr = 0.05 and various jet velocities.  At higher concentrations (Cm ≥ 1.0%), the jet velocity had less influence on mixing. Figure 5.11 plots the modified mixing index at various velocity ratios for a 2.0% softwood suspension at Up = 0.5 m/s and virtually identical impeller speed.  119  The degree of mixing clearly improved with the addition of the mechanical mixer. However, mixing improved only slightly with increasing jet velocity since the fibre networks were still robust and energy dissipation from jet impaction was insufficient to disrupt the fibre networks. The mixing quality was worse when the jet reached the far wall of the pipe ahead of the impeller at R = 12.2, likely due to the robust fibre networks causing the jet to adhere to the far wall downstream, behaviour similar to that for a tee alone. Tomographic images comparing tracer distribution for the same operating conditions are presented in Figure 5.12. For a jet reaching the far wall of the pipe at R = 12.2, the high-conductivity regions were concentrated at the upper part of the images and persisted as a clump downstream, as illustrated in Figure 5.12a. At R = 16.5, the jet impinged on the far wall and recirculated to the core of the pipe. Energy dissipation from the jet impingement and rotating impeller likely caused less non-uniformity, as shown by more spreading of high-conductivity regions in the cross-sections downstream of the P2 plane in Figure 5.12b. Although a higher velocity ratio of 24.1 provided slightly better mixing than R = 8.1 for a jet penetrating to the core of the pipe, this condition may be undesirable from a practical point of view since impingement on the opposite pipe wall creates significant stress there. The design of an in-line mechanical mixer should therefore be based on the velocity ratio for the jet penetrating to the axis of the pipe, as summarized in Table 4.3 of Chapter 4.  120  100  M' (%)  R  N (rpm)  9.62 8.14 12.2 16.5 24.1  NA 408 403 410 418  P7  P8  10  P2  P3  P4  P5  P6  Impeller 1 0  5  10  15  20  25  x/D  Figure 5.11: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Cm = 2.0%, Up = 0.5 m/s, Dr = 0.05, almost constant impeller rotation speeds and various jet velocities.  5.3.4 Effect of impeller rotational speed on mixing quality in pulp suspensions Figure 5.13 shows the effect of impeller speed on the mixing quality for a hardwood pulp suspension with Cm = 1.0%. The mixing was poor for a tee mixer alone, but improved substantially in the presence of the impeller. At a lower mainstream velocity (Up = 0.5 m/s), long residence time in the high-shear zone provided profoundly improved mixing with increasing impeller rotational speed, as shown in Figure 5.13a. Decay of turbulence likely occurred, accompanied by reflocculation downstream, as clearly shown by worse mixing downstream of P4 for  121  (a)  Top  (b)  Figure 5.12: Tomographic images for softwood pulp suspension at Up = 0.5 m/s, Cm = 2.0%, Dr = 0.05 with (a) R = 12.2, N = 403 rpm and (b) R = 16.5, N = 410 rpm. Locations of planes P1 to P8 are shown in Figure 3.5.  122  pulp suspension than for water at N = 800 rpm. The crowding number of 27.6 for the 1.0% hardwood pulp suspension exceeded 16, the gel crowding number, suggesting that fibres interacted and began to flocculate when the shear stress was lower than the suspension yield stress downstream of the impeller. For this velocity, the quality of mixing with the perpendicularly orientated static mixer was lower than that for all non-zero rotational speeds investigated. The suspension flowed through the gaps between the impeller and inner pipe wall at relatively low velocity, and the turbulence created by the static mixer was significantly less than that provided by the mechanical mixer. At a higher mainstream velocity (Up = 2.0 m/s), the residence time was low enough that the impeller speed had little influence on the mixing quality. An increase in rotational speed, however, slightly enhanced downstream mixing as shown in Figure 5.13b. The mixing quality was similar for the first three planes (P2 – P4), and then became slightly better with increasing impeller speed. A comparison of suspension mixing at N = 400 and 800 rpm with water (dashed lines) at the same impeller speeds showed that turbulence was achieved in the high-shear zone for both impeller speeds, but reflocculation likely occurred faster downstream at N = 400 rpm (downstream of P5) than N = 800 rpm (downstream of P7), due to reduced downstream turbulence. Figure 5.14 shows the influence of rotation speed on mixing for a long-fibre softwood pulp suspension with a higher Cm of 3.0%. At Up = 1.0 m/s, mixing with the parallel static impeller was poor, not significantly better than for the tee mixer alone,  123  (a) 100  Cm = 1.0%  M' (%)  10  P2  P3  P4  P5  P6  P7  P8  N (rpm)  R  NA 12.46 0 (//) 11.92 0 ( ) 11.92 406 12.46 603 11.85 802 11.97 804 (w) 11.51  1  0.1  Water  Impeller  0  5  10  15  20  25  x/D  (b) 100  N (rpm) NA 407 603 802 400 (w) 801 (w)  M' (%)  10  1  R 1.87 2.07 2.06 2.07 2.06 2.08  Impeller water 0.1  0  5  10  15  20  25  x/D  Figure 5.13: Modified mixing index as a function of dimensionless distance downstream of injection for hardwood pulp suspension with Cm = 1.0% and water (w), Dr = 0.05 and almost constant velocity ratios with various impeller speeds at: (a) Up = 0.5 m/s (b) Up = 2.0 m/s.  124  as shown in Figure 5.14a. At Up = 1.0 m/s, the mixing quality improved with increasing impeller speed, predominantly occurring in the high-shear (P2) zone, with lack of downstream mixing compared to a lower-concentration, shorter-fibre hardwood suspension and water, since strong fibre networks re-established immediately downstream of the impeller. The energy from the impeller and the flow (main stream and jet) were unable to disrupt the networks in the high-shear zone, even at an impeller speed as high as 800 rpm, as shown by considerably worse mixing at P2 than for water at a lower impeller speed (N = 400 rpm). Further increasing the impeller speed likely improves mixing and provides a similar level of turbulence as for water. However, this probably occurs only in the high-shear zone without improved downstream mixing. At this mainstream velocity, the highconcentration suspension flow, carrying high momentum, provided high energy dissipation when it impinged on the perpendicularly-orientated static impeller, with the mixing quality similar to that of the mechanical mixer at N = 800 rpm. For water flow at the identical mainstream velocity, however, the perpendicular configuration only matched the mixing provided by the mechanical mixer at N = 400 rpm, as shown in Figure 5.3. At a higher mainstream velocity (Up = 2.0 m/s), the impeller improved mixing significantly, as shown in Figure 5.14b, with slightly enhanced downstream mixing compared to a lower mainstream velocity. The impeller speed, however, had little effect on mixing, likely due to shorter residence time in the highshear zone.  125  (a) 100  Cm = 3.0% P2  10  P3  P4  P5  P6  P7  P8  N (rpm) R     M' (%)  NA 6.57 0 (//) 6.24 0 ( ) 6.09 405 6.14 605 6.17 802 6.12 401(w) 6.12  1  Cm = 0% (water) Impeller  0.1  0  5  10  15  20  25  x/D  (b) 100  Cm = 3.0% N (rpm)  M' (%)  10  Cm = 0% (water)  1  R  NA 2.06 403 2.05 601 2.06 808 2.05 400(w) 2.06  Impeller 0.1  0  5  10  15  20  25  x/D  Figure 5.14: Modified mixing index as a function of dimensionless distance downstream of injection for softwood pulp suspension with Cm = 3.0% and water (w), Dr = 0.05 and almost identical velocity ratios with various impeller speeds at: (a) Up = 1.0 m/s (b) Up = 2.0 m/s.  126  Figure 5.15 illustrates the effect of impeller speed and mainstream velocity on mixing quality for a softwood pulp suspension with Cm = 3.0%. For N ≈ 400 rpm, mixing at Up = 1.0 m/s was considerably lower than at Up = 3.0 m/s, both in the high-shear (P2) zone and downstream. Increasing impeller speed to 800 rpm for Up = 1.0 m/s improved mixing in the high-shear zone, approaching that for Up = 3.0 m/s, with a lack of downstream mixing. Since mixing at Up = 3.0 m/s does not improve much with increasing impeller speed, due to short residence time in the high-shear zone, a further increase in shear at Up = 1.0 m/s likely provides better mixing than at Up = 3.0 m/s in the high-shear zone, with no improvement in mixing downstream. The downstream mixing depends on initial mixing in the high-shear zone, and is likely better than at Up = 3.0 m/s for very high impeller speeds.  100  P2  P3  P4  P5  P7  P8  Up(m/s)  R  N (rpm)  1.0 1.0 3.0  4.27 4.24 4.13  411 808 440  P6  M' (%)  10  1  0.1 0  5  10  15  20  25  x/D  Figure 5.15: Modified mixing index as a function of dimensionless distance downstream of injection for softwood pulp suspension with Cm = 3.0%, Dr = 0.05 and similar velocity ratio for various impeller speeds and mainstream velocities.  127  5.3.5 Effect of fibre mass concentration on mixing quality Figure 5.16 plots the modified mixing index for various fibre mass concentrations of a softwood pulp suspension at Up = 1.0 m/s, N = 400 rpm, nearly identical jet-to-pipe velocity ratios and identical diameter ratio. The mixing quality improved significantly as the consistency decreased. Fibre mass concentration strongly influenced the mixing when turbulent flow could not be achieved.  100  Cm = 3.0%  10  M' (%)  Cm = 2.0% Cm = 1.0%  P2 1  P3 Impeller  Cm = 0.5%  P4 P5  Water  P6  0.1 0  5  10  15  P7 20  P8 25  x/D  Figure 5.16: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Up = 1.0 m/s, R = 6, N = 400 rpm, Dr = 0.05 for various fibre mass concentrations.  Figure 5.17 shows the influence of consistency on the mixing quality for shorter-fibre hardwood suspension at a higher impeller speed, 800 rpm. The mixing quality improved profoundly as Cm decreased from 2.0 to 1.0%, at which the imposed shear was sufficient to disrupt the plug and provide turbulence. The mass  128  concentration did not, however, have a strong effect on the mixing quality for Cm ≤ 1.0% since the suspension flow was turbulent, with mixing quality similar to that for water.  100  10  M' (%)  Cm(%) 0 0.5 1.0 2.0 3.0  P2  1  P3 P4  0.1  P5  Impeller  P6  P7  P8  0.01 0  5  10  15  20  25  x/D  Figure 5.17: Modified mixing index as a function of dimensionless distance downstream for hardwood pulp suspension at Up = 1.0 m/s, R = 6, N = 800 rpm, Dr = 0.05 for various fibre mass concentrations.  For Cm < 1.0%, the mixing quality was similar for the first three planes (P2 – P4 in Figure 1) downstream of the injection. Mixing for the hardwood pulp suspension with Cm = 1.0% was worse downstream than for Cm = 0 and 0.5%, likely due to reduction of turbulence accompanied by reflocculation, occurring since the energy required to maintain turbulence was not sustained at the higher mass concentration. The crowding number of 13.8 for Cm = 0.5% was less than 16, indicating that fibres were free to move relative to one another, so that the suspension was essentially dilute. On the other hand, the crowding number for the  129  1.0% hardwood pulp suspension exceeded the gel crowding number of 16. Accordingly, reflocculation likely occurred only at Cm = 1.0%. The mixing quality with the mechanical mixer for a hardwood suspension flow at Cm = 0.5% was similar to that for water without fibres, whereas mixing with a tee mixer and turbulent flow was better for a dilute hardwood suspension than for water, as discussed in section 4.3.7. For tee mixing alone, mixing is achieved by turbulent shear. The shorter and smaller diameter fibres could alter the turbulence structure in the bulk, carrying turbulent eddies with them and colliding with each other, thereby promoting turbulent dispersion. For the mechanical mixer, however, the energy dissipated near the rotating impeller probably dominated, modifying the turbulence structure in the bulk and leading to similar mixing quality for water and a dilute hardwood suspension. Tomograhic images showing turbulent and plug flow for hardwood suspensions are presented in Figure 5.18. For Cm = 0.5%, the flow was turbulent and the tracer distributed uniformly downstream of P2, as illustrated by uniform green in the cross-sections, similar to that upstream of tracer injection (P1) in Figure 5.18a, indicating efficient mixing downstream of the impeller. Figure 5.18b shows the low-conductivity region in blue downstream of P2, likely a plug caused by fibre reflocculation at a higher fibre mass concentration, 3.0%. This region occurred immediately downstream of P2, indicating rapid reflocculation as the local rate of energy dissipation decreased downstream.  130  (a)  Top  (b)  Figure 5.18: Tomographic images for hardwood pulp suspension at Up = 1.0 m/s, R = 6, N = 800 rpm, Dr = 0.05 with (a) Cm = 0.5% and (b) Cm = 3.0%. Locations of planes P1 to P8 are shown in Figure 3.5.  131  Figure 5.19 illustrates the influence of fibre mass concentration on the mixing quality for a softwood suspension at a higher mainstream velocity of 3.0 m/s and virtually identical impeller speeds. For Cm ≤ 1.0%, the energy provided by the impeller and the flow was sufficient to disintegrate the fibre networks and provide turbulence in the high-shear zone. Reflocculation and decaying turbulence then probably occurred rapidly for Cm = 1.0%, as shown by the mixing quality being worse downstream of the P3 sensor plane compared to water and 0.5% pulp suspension in the turbulent flow regime. At higher mass concentrations (Cm ≥ 2.0%), on the other hand, the energy supplied was insufficient to provide the same level of turbulence as for water, even in the high-shear zone immediately downstream of the impeller. Reflocculation also likely occurred more rapidly than for the 1.0% pulp suspension, with significantly lower mixing quality downstream.  100  Cm(%) N (rpm) 0 0.5 1.0 2.0 3.0  M' (%)  10  401 404 402 433 440  P2 1  P3  P4  P5  Impeller  P6  P7  P8  0.1 0  5  10  15  20  25  x/D  Figure 5.19: Modified mixing index as a function of dimensionless distance downstream for softwood pulp suspension at Up = 3.0 m/s, R = 4, Dr = 0.05 and virtually identical rotational speeds for various fibre mass concentrations.  132  5.3.6 Hardwood vs. softwood fibres Figure 5.20 shows that the mixing quality for the short-fibre hardwood pulp suspension was considerably better than for the softwood suspension at Cm = 1.0%, Up = 0.5 m/s and N = 400 rpm. This was likely due to less robust fibre networks. However, the imposed shear was not sufficient to disintegrate the fibre networks and provide turbulence in the suspension, as shown from the mixing quality being significantly worse than for water in turbulent flow.  100 Softwood fibers Hardwood fibers Water (no fibers)  M' (%)  10  P2 P3  1  P4 P5  Impeller  P6  P7  P8  0.1 0  5  10  15  20  25  x/D  Figure 5.20: Modified mixing index as a function of dimensionless distance downstream for softwood and hardwood pulp suspensions at Cm = 1.0% Up = 0.5 m/s, R = 12, N = 400 rpm.  The influence of fibre type on the degree of mixing for a lower mass concentration of 0.5% and a higher main stream velocity of 2.0 m/s is presented in Figure 5.21. In this case, the fibre properties did not have such a strong influence  133  on mixing quality since the turbulence was generated in both softwood and hardwood suspensions, with mixing quality approaching that for water.  100 Softwood fibers Hardwood fibers Water (no fibers)  M' (%)  10  P2 1  P3 P4  P5  Impeller  P6  P7  P8  0.1 0  5  10  15  20  25  x/D  Figure 5.21: Modified mixing index as a function of dimensionless distance downstream for softwood and hardwood pulp suspensions with Cm = 0.5% Up = 2.0 m/s, R = 2 and N = 400 rpm.  Table 5.1 summarizes the influences of jet penetration, flow regime and impeller speed on mixing quality with an in-line mechanical mixer. For water, a jet penetrating to the centre of the pipe clearly provided better mixing than the jet attaching to the pipe wall, and mixing was even better when the jet impinged on the far wall at higher jet-to-pipe velocity ratio. A similar effect of jet penetration on mixing was observed for pulp suspensions, with considerably worse mixing when the jet reached the far wall of the pipe for the plug flow regime, likely due to the reestablishment of robust fibre networks downstream of the impeller causing the jet to adhere to the wall. Jet impingement on the far wall, however, did not provide  134  significantly better mixing than that for jet penetration to the core of the pipe, likely because the energy input was insufficient to disrupt the fibre networks for both cases. The flow regime had a strong influence on mixing in pulp suspensions. For similar jet penetration, mixing worsened substantially as the flow regime changed from turbulent to mixed and plug flow. For the mixed flow regime, increasing the impeller speed clearly enhanced mixing in both the high-shear (P2) zone and downstream, whereas mixing improved predominantly in the high-shear zone for the plug flow. Supplementary data relevant to this chapter are provided in Appendix B.  5.4 Conclusions For water in turbulent flow, the quality of mixing with perpendicular static and low-speed impellers was almost independent of the mainstream velocity for identical jet-to-pipe velocity and diameter ratios, since the residence time in the high-shear zone had little influence. At higher impeller speeds (N ≥ 600 rpm), the residence time in the high-shear zone became significant, with mixing improving considerably as the mainstream velocity decreased. Mixing improved substantially with increasing impeller speed at a low mainstream velocity (1.0 m/s), likely due to increased residence time in the high-shear zone, whereas the rotation speed had less effect at higher mainstream velocities (Up ≥ 2.0 m/s). At low impeller speed (400 rpm), the mixing quality improved substantially with increasing velocity ratio as  135  Table 5.1: Summary of the influences of jet penetration, flow regime and impeller speed on mixing Jet-toSuspension Impeller Mainstream pipe Parameter concentration, speed, N velocity, Up velocity investigated (m/s) Cm (%) (rpm) ratio, R  Jet penetration in water  Jet penetration in pulp suspension  Flow regime  0 (water)  2.0 (SW)  0.5 (SW) 0.5 (SW) 0.5 (SW)  Impeller speed (plug flow)  1.0 (HW)  3.0 (SW)  Modified mixing index, M' Jet penetration x/D = x/D = x/D= 2.41a 12.2 22.1b Attaching to near wall 4.41 0.87 0.51  2.0, 3.0  2.06  1.0, 2.0  6.15  0.5, 1.0  12.2  0.5  24.6  408  0.5  8.14  403  0.5  12.2  418  0.5  24.1  1.0  3.99  2.0  4.13  Turbulent Almost centre 2.34 0.92 0.58  3.0  4.13  Turbulent Almost centre 2.84 0.87 0.44  400  400  2.0 (SW) Impeller speed (mixed flow)  Flow regime  1.0  6.14  No impeller  1.0  12.5  406  1.0  12.5  802 No impeller  1.0  12.0  1.0  6.35  406  1.0  6.14  802  1.0  6.12  a  High-shear zone (P2)  b  Furthest distance downstream (P8)  Centre 3.02 0.4 0.21 Turbulent Reaching far wall 3.32 0.62 0.25 Impinging on far wall 1.43 0.12 0.06  Plug  Almost centre 6.96 6.87 6.61 Reaching far wall 8.88 9.17 9.75 Impinging on far wall 7.59 6.67 6.02  Mixed  Almost centre 2.5 1.49 1.17  Plug  Mixed  Plug  Almost centre Reaching far wall Reaching far wall Reaching far wall  5.58 4.45 4.36 5.79 5.56 5.1 4.04 1.47 1.26 2.41 0.26 0.16  Almost centre 8.89 9.24 9.55 Almost centre 7.78 7.26 7.42 Almost centre 4.95 4.89 4.69  the mixing mode changed from wall-source to jet-mixing and jet-impaction. However, the mixing quality was worse when the jet stream reached the far wall of the pipe upstream of the impeller at a jet-to-pipe velocity ratio of 12.2.  136  For pulp fibre suspensions, mixing depended strongly on the flow regime. Mixing downstream improved profoundly when the plug was disrupted, approaching that for water when the flow became turbulent. With addition of an impeller, the turbulent flow regime was obtained for 0.5% softwood pulp suspension at a mainstream velocity of 2.0 m/s, whereas 4.0 m/s was required without the impeller. At a higher mass concentration (Cm = 2.0%), the fibre plug was disrupted and mixing improved profoundly with increasing mainstream velocity, but downstream mixing was poor, likely due to rapid reflocculation. For a dilute softwood pulp suspension (Cm = 0.5%), mixing improved with increasing jet velocity, whereas the jet velocity had little influence at Cm = 2.0%. Mixing quality worsened when the jet penetrated to the far wall of the pipe for all mass concentrations investigated. For a dilute hardwood pulp suspension with Cm = 1.0%, mixing was improved by the addition of an impeller. A stationary impeller provided better mixing when perpendicular than when parallel to the flow. As the rotational speed increased, mixing improved significantly, both in the high-shear zone and downstream, for a low mainstream velocity (Up = 1.0 m/s), whereas it was only slightly better at Up = 2.0 m/s, likely due to less residence time in the high-shear zone. At a higher mass concentration of 3.0% for long-fibre softwood pulp suspension, enhancement of mixing with increasing rotation speed occurred mainly in the high-shear zone, with almost negligible improvement downstream. Fibre mass concentration and properties profoundly affected mixing when the supply of energy was insufficient to provide turbulence. Mixing improved substantially with decreasing consistency, and was better for shorter hardwood pulp  137  fibres than for softwood fibres. The mass concentration and fibre properties did not, however, strongly influence the mixing quality when the flow was turbulent, since the turbulence disrupted the plug, with mixing quality approaching that for water. Unlike tee mixing, the quality of mixing with an in-line mechanical mixer for a dilute hardwood suspension flow was not significantly better than for water.  138  6. Gas Dispersion in Horizontal Pulp-FibreSuspension Flow7  6.1 Introduction Effective contacting of a gas phase with pulp and its uniformity in suspension are essential for three-phase (gas-water-fibre) systems in various pulping operations such as oxygen delignification, oxidative extraction and ozone bleaching. Good mixing provides the uniform dispersion of the gas phase and high interfacial area between the phases, promoting mass transfer and hence overall reaction. Predistribution of the gas phase in pulp suspensions is usually achieved by in-line mixers ahead of tower reactors. Gas is injected into pulp suspensions ahead of, or inside, various mixers, including peg and high-shear mixers, and the mixture then flows along the pipes before entering the tower reactors. In spite of its importance, gas-suspension contacting is not well characterized, with literature on gas dispersion in pulp suspension flow in horizontal pipes being rare. Gas-liquid two-phase horizontal flow also occurs in various industrial applications. The distribution of gas in this flow depends on several factors including 7  This chapter is very similar to parts of two papers accepted for publication: Yenjaichon, W., Grace,  J.R., Lim, C.J., Bennington, C.P.J., 2012. Characterisation of gas mixing in water and pulpsuspension flow based on electrical resistance tomography, Chem. Eng. J.; and Yenjaichon, W., Grace, J.R., Lim, C.J., Bennington, C.P.J., 2012. Gas dispersion in horizontal pulp-fibre-suspension flow, Int. J. Multiphase Flow.  139  gas-to-liquid flow ratio, fluid properties and pipe geometry. Several flow regimes have been identified (e.g. Govier and Aziz, 1972; Mandhane et al., 1974;Taitel and Dukler, 1976; Barnea, 1987); most authors agree on six flow patterns: stratified flow, wave flow, slug flow, annular flow, elongated bubble (or plug) flow and bubble flow. Some (Barnea, 1987; Andreussi et al., 1999; Ekambara et al., 2008) divide bubble flow into two sub-flow patterns, including dispersed bubble flow and bubble flow, even though the distinction between them is not always clear. At high superficial liquid velocities and low superficial gas velocities, gas forms the dispersed phase, whereas liquid constitutes the continuous phase. Gas is uniformly dispersed as small bubbles throughout the continuous liquid phase for dispersed bubble flow. Buoyancy causes the bubbles to migrate to the top of the pipe at relatively low liquid velocities, leading to asymmetric distribution of the bubbles and changes in their size and shape, referred to as the bubble flow regime. Andreussi et al. (1999) classified these two sub-flow patterns by the bubble shape, i.e. spherical or nearly-spherical bubbles for dispersed bubble flow and ellipsoidal or pulsating shapes for bubble flow. At intermediate superficial liquid velocities and low gas velocities, discrete elongated bubbles slide along the inner top surface of the pipe. This is known as elongated bubble flow. At higher gas velocities, the bubbles become longer, occupying a larger portion of the pipe and separated by slugs of liquid containing entrained small bubbles. This is defined as slug flow. At low superficial liquid and gas velocities, the gas phase occupies the upper portion of the pipe, while liquid flows at the bottom, with a smooth gas-liquid interface. This is referred to as  140  stratified flow. At higher gas velocities, waves are produced at the interface, and this is called the wave flow pattern. Further increasing gas velocity causes a liquid annulus to form around the gas core, denoted an annular flow. In this chapter, four flow regimes − stratified flow (S), elongated bubble flow (EB), bubble flow (B) and dispersed bubble flow (DB) − were examined. Since the gas and liquid phases are fully segregated at the top and bottom of the pipe, respectively, for both stratified and wave flow patterns, both of these flow patterns are referred to as stratified flow in this study. The distribution of the dispersed phase and interfacial area between the phases are key aspects for evaluating gas-liquid flow, generally determined from the distribution of local void fraction and bubble size. Measurement techniques utilized to measure local void fraction and bubble diameter in air-water horizontal flow are summarized in Table 6.1. Conductance probes, resistivity probes and hotfilm anemometers can measure local void fractions and velocities, but they interfere with the process. Measuring local void fractions throughout a pipe cross-section with these techniques is also tedious, since many measurements are required to cover the cross-section. High-speed photography is non-invasive and provides high spatial and temporal resolution, but it is not suitable for opaque pipes and fluids such as pulp fibre suspensions. Magnetic resonance imaging is also non-invasive, providing three-dimensional velocity and concentration data. However, fluids or pipes need to be non-conductive or non-magnetic to avoid interactions with the magnetic field, and this technique is expensive.  141  Table 6.1: Summary of techniques for measuring void fraction and bubble size for different flow patterns in air-water horizontal flow (DB = dispersed bubble, SL = slug, S = stratified, W = wavy) Technique  Dubblesensor resistivity probe  Conductance probe  MRI  Dp (mm)  Usg (m/s)  Flow Pattern  50.3  3.74 5.71  0.25 1.37  DB  50.3  3.74 6.59  0.21 1.34  DB  50.3  3.74 6.59  0.21 1.34  DB  53 and 90  0.6 3.5  0.5 20  SL  31 and 53  0.6 3.5  0.5 20  SL  18 and 50  4.40 4.53  0.08 1.22  DB  25.4  0.08 1.3  0.3 14  S, W, SL  50.3  3.8 5.0  0.25 0.8  DB  50.3  1.1 2.2  0.27 2.2  SL  25.4  1.3 2.9  102.6  2.1 4.7  Hot-film anemometry  High-speed camera  a  Usl (m/s)  a  DB  a  DB  Measured parameters Local void fraction,interfacial area concentration, bubble interface velocity, bubble diameter Void fraction, interfacial area concentration, bubble diameter Local void fraction, interfacial area concentration, bubble interface velocity, bubble frequency Slug length, void fraction, bubble velocity Local void fraction, slug length, slug frequency, bubble diameter Local void fraction, bubble diameter Void fraction, interfacial area concentration Local void fraction, bubble frequency, turbulent intensity, mean liquid velocity Local void fraction, turbulent intensity, mean liquid velocity Bubble diameter, bubble size distribution Bubble diameter, bubble size distribution  Reference  Kocamustafaogullari and Wang (1991)  Kocamustafaogullari et al. (1994)  Kocamustafaogullari and Huang (1994)  Nydal et al. (1992) Andreussi et al. (1993) Andreussi et al. (1999) Reyes et al. (1998) Iskandrani and Kojasoy (2001) Lewis et al. (2002) Razzaque et al. (2003) Sanders et al. (2004)  Authors did not provide Usg, but reported air volume fractions of 0.0007 to 0.003 in both cases.  142  Electrical resistance tomography (ERT) has some advantages over other techniques since it is non-intrusive, inexpensive and suitable for opaque systems. ERT has been used extensively to measure gas holdup in various systems (e.g. Jin et al., 2007; Razzak et al., 2007; Abdullah et al., 2011), but few data are available in the literature for gas holdup in horizontal pipes (Dong et al., 2003; Meng et al., 2010). A limitation of ERT for gas-liquid horizontal flow is that some electrodes lose electrical contact when the gas concentrates at the top of the pipe. Ma et al. (2001) applied an alternative voltage data collection method when gas occupied the upper part of the pipe by detecting the water level in the cross-section, and used conventional ERT data collection and image reconstruction for the bubble flow pattern. In this study, resistor adaptors provided an alternate current path when the gas phase hindered the electrodes, enhancing voltage measurement and image reconstruction. To the best of the author’s knowledge, no previous ERT studies of gas-liquid horizontal flow using resistor adaptors have been reported in the literature. Mixing quality can be measured according to the uniformity and size of the dispersed gas phase to describe gas hydrodynamics in the system. One of the most comprehensive methods was proposed by Danckwerts (1952). Two parameters characterized the mixing quality based on the statistical theory of turbulence, with details provided in section 2.5. In pulp fibre suspensions, Is (or modified forms of Is) have been applied extensively to measure mixing quality (e.g. Kamal and Bennington, 2000; Kourunen et at., 2011). However, few studies have used Ls to evaluate the degree of mixing (Paterson and Kerekes, 1985, 1986) due to the  143  extensive effort needed to collect the data for evaluation of the coefficient of correlation. Since it covers the full pipe cross-section, ERT can be utilized to quantify the degree of mixing based on the coefficient of variation, with a number of applications discussed in section 2.8.3. The application of ERT to measure Ls provides extensive analysis, without excessive work to determine the coefficient of correlation. In this chapter, a gas mixing index, based on variation of local gas holdup in a crosssectional plane, and the scale of segregation are applied to investigate gas uniformity and the relative size of gaseous entities such as bubbles for various flow conditions, in order to provide better understanding of gas dispersion in pulp suspension flows in horizontal pipes. This chapter focuses on developing flow of airwater/suspension mixing downstream of a 90 tee.  6.2 Experimental Details Details of the experimental setup are described in section 3.1. The air injection tube diameter was 15.9 mm (Dr = 0.208) for all the experiments in this chapter, unless identified. The experiments were conducted over a range of fibre mass concentrations (0 – 3.0%), superficial liquid/pulp velocities (0.5 – 5.0 m/s) and superficial gas velocities (0.11 – 0.44 m/s), using the procedure explained in section 3.2.2. The uniformity of gas in cross-sectional planes along the pipe was quantified by the gas mixing index, described in section 3.3.2.  144  6.3 Results and Discussion 6.3.1 ERT for evaluating gas-liquid horizontal flow Figure 6.1 compares tomographic images for Usl = 0.5 m/s and Usg = 0.11 m/s obtained with and without resistor adaptors. Visual observation showed clearly that stratified flow occurred for these flow conditions. The dashed line shows the vertical bisector, rotated due to misalignment between the electrode position in the image reconstruction and the actual electrode position. The images show the distribution of local gas holdup, with the red and yellow regions representing air, injected between planes 1 and 2, and blue representing water. The gas migrated rapidly to the top of the pipe, resulting in stratified flow. The contact of gas with the electrodes prevented electrical current flow through the pipe cross-section, leading to anomalous voltage data and reconstructed images, as illustrated by blue at the top of the pipe in Figure 6.1a. Tomographic images obtained with the resistor adaptors in Figure 6.1b, on the other hand, show high gas-concentration regions at the top of the image, consistent with stratified flow. The images are imperfect, however, as they show colours from blue (low gas concentration) at the outside to red (high gas concentration) at the centre of the gas region, in contrast to a stratified flow in which only gas occupied the top of the pipe. Electrical current flows through the adaptors when the gas phase hinders electrodes from injecting the current, and voltages are measured across the adaptors. The images are then reconstructed from these voltages, with water as the reference medium, leading to lowconcentration colour regions around the outer area (Talideh, 2012). Those regions  145  were not due to water wetting the electrodes, since measurements with a dry pipe (no water) still resulted in colour layers. This is referred to as the “edge effect” in this study. (a)  Top  (b)  Figure 6.1: Tomographic images for air-water flow at Usl = 0.5 m/s, Usg = 0.11 m/s, Dr = 0.208: (a) without resistor adaptors (b) with resistor adaptors. Locations of planes P1 to P8 are shown in Figure 3.2.  146  The average gas holdup from the ERT measurements is compared in Table 6.2 to that determined from visual observation at x/D = 22.1 (plane 8) for several superficial gas velocities in stratified flow. The gas holdup measured with the adaptors was close to that from visual observation for all flow conditions, whereas that without the adaptors was much lower than observed visually, likely due to the presence of a virtually gas-free liquid phase at the top of the image caused by false voltage data. While there are limitations of measurements with the resistor adaptors, the results with the adaptors were clearly better than those without them. The adaptors were therefore used in all subsequent tests.  Table 6.2: Comparison of gas holdup measured without resistor adaptors, with adaptors and visual observation for the stratified flow pattern Usl (m/s) 0.5  Usg (m/s) 0.11 0.22 0.33 0.44  Average gas holdup at x/D = 22.1 without adaptors with adaptors visual observation 0.124 0.331 0.31 0.135 0.336 0.32 0.163 0.339 0.32 0.195 0.348 0.33  6.3.2 Gas mixing indices 6.3.2.1 Gas mixing index for air-water flow Figure 6.2 illustrates the influence of superficial gas velocity on the mixing quality for dispersed bubble flow at Usl = 5.0 m/s. At the lowest gas velocity, 0.11 m/s, the gas mixing index (defined in equation 3.12) decreased, i.e. the mixing quality improved, along the pipe, with air injected at the bottom of the pipe just  147  upstream of P2. The bubbles were concentrated at the bottom of the pipe at P2 and P3, then migrated to the centre of the pipe and dispersed throughout the pipe crosssection downstream, resulting in better mixing, as shown by the displacement of high gas-concentration regions (green and yellow) along the pipe in Figure 6.3a. At higher superficial gas velocities, the mixing quality was better downstream from P2 to P4, with air bubbles travelling from the bottom to the core of the pipe. The bubbles then accumulated and formed bigger ones, rising to the top of the pipe downstream due to buoyancy, leading to significant demixing. Tomographic images for Usg = 0.44 m/s in Figure 6.3b show that bubbles congregated and migrated to the top of the pipe downstream. The mixing quality decreased with increasing superficial gas velocity, likely due to the presence of more and larger bubbles. Note that the colour scale indicating gas holdup differs for each figure in which tomographic images are presented.  100  Usg(m/s)  Mg (%)  0.11 0.22 0.33 0.44  10  P2 P3  4  0  5  P4  P5  10  P6 15  P8  P7 20  25  x/D  Figure 6.2: Gas mixing index as a function of dimensionless distance downstream for dispersed bubbles in air-water flow at Usl = 5.0 m/s and Dr = 0.208 for various superficial gas velocities. Locations of planes P2 to P8 are shown in Figure 3.2.  148  (a)  Top  (b)  Figure 6.3: Tomographic images for air-water flow at Usl = 5.0 m/s and Dr = 0.208 for: (a) Usg = 0.11 m/s (b) Usg = 0.44 m/s. Colour scales indicate gas holdup. Locations of planes P1 to P8 are shown in Figure 3.2.  149  Figure 6.4 illustrates the effect of flow pattern and superficial gas velocity on mixing for air-water flow at Usl = 2.0 m/s. The flow patterns were dispersed bubble flow, bubble flow, elongated bubble flow or transition between the last two, depending on the distance downstream. At Usg = 0.11 m/s, the bubbles migrated from the bottom of the pipe (P2) to the centre (P3), with similar bubble size, and hence virtually the same mixing quality. The bubbles dispersed in the cross-section at P3, and then coalesced and rose rapidly to the upper part of the pipe at P4, with considerably larger bubbles than at Usl = 5.0 m/s recorded by the high-speed camera. The gas mixing index also shows that the mixing quality at the P4 plane was significantly worse than for the two preceding planes. Bubbles continued to coalesce along the pipe and finally became elongated bubbles at P7 and P8. Clearly, the mixing quality was worse downstream with the change from dispersed bubble to bubble and elongated bubble flow. At higher superficial gas velocities, the flow pattern at each downstream position was similar to that at Usg = 0.11 m/s. However, gas dispersion in the main stream was less uniform since the larger bubbles occupied the top of the pipe. Figure 6.5 shows the influence of superficial gas velocity on mixing for stratified flow at Usl = 0.5 m/s. For Usg = 0.11 m/s, the mixing quality deteriorated downstream due to changes in flow pattern. Bubbles migrated rapidly from the injection point at the bottom of the pipe, reaching the top at P2. They then disintegrated into smaller bubbles, moving rapidly and occupying the top of the pipe at P3. The gas segregated, leading to stratified flow at P4, with considerably worse mixing. Mg then increased slightly and reached a plateau downstream, consistent  150  with visual observations showing that fully-developed stratified flow was approached. At higher superficial gas velocities, the flow pattern along the pipe was similar to that for Usg = 0.11 m/s. However, the gas mixing index was higher for higher gas velocities when the air was introduced to the main stream, since more bubbles were present, with less uniform dispersion over the cross-section. The gas mixing index, however, levelled off and reached approximately the same value of ~ 0.7, independent of the superficial gas velocity when stratified flow was approached. One would expect a value closer to 1.0 since only gas occupied the top of the pipe. The discrepancy was due to the spreading of the gas-liquid interface (i.e. the edge effect from the image reconstruction process) reducing the standard deviation (and hence Mg), whereas spreading of the interface did not cause significant error in the mean gas holdup (Table 6.2), which only relies on averaging.  100 DB - Dispersed bubble flow EB - Elongated bubble flow B - Bubble flow P7 P6 P5  80 60  P4  Mg (%)  40 P2  P3 B/EB  B/EB  B  20 B 10  P8  0  EB  Usg(m/s) 0.11 0.22 0.33 0.44  DB  5  EB  10  15  20  25  x/D  Figure 6.4: Gas mixing index as a function of dimensionless distance downstream for air-water flow at Usl = 2.0 m/s and Dr = 0.208 for various superficial gas velocities.  151  100 90 80  Stratified flow  Mg (%)  70 P4  60  P6  P5  Usg(m/s)  P3  0.11 0.22 0.33 0.44  P2  50  P8  P7  40 0  5  10  15  20  25  x/D  Figure 6.5: Gas mixing index as a function of dimensionless distance downstream for the stratified air-water flow at Usl = 0.5 m/s and Dr = 0.208 for various superficial gas velocities.  Figure 6.6 illustrates the effect of superficial liquid velocity on mixing for Usg = 0.11 m/s. The mixing quality depended greatly on the flow pattern. Worse mixing was observed with decreasing superficial liquid velocity as the flow pattern changed from dispersed bubble to bubble, elongated bubble and finally stratified flow. At Usl = 5.0 m/s, the turbulent shear in the liquid overcame the buoyancy forces on the bubbles, dispersing bubbles throughout the pipe cross-section downstream. At Usl = 4.0 m/s, the bubbles dispersed throughout the cross-section downstream of P3. They then congregated and migrated to the top of the pipe downstream of P5, with significantly lower mixing quality. At lower liquid velocities, the bubbles coalesced rapidly and occupied the top of the pipe, becoming elongated bubbles, with considerably worse mixing. The mixing deteriorated at Usl = 0.5 m/s as stratified  152  flow was approached. The flow regimes in this study cannot be compared directly with air-water flow regimes in the literature as those are for fully-developed flow, but our results at x/D = 22.1, the furthest position downstream, appear to agree relatively well with flow regime maps in the literature (Govier and Aziz, 1972; Mandhane et al., 1974; Taitel and Dukler, 1976), e.g. elongated bubble flow at Usl = 1.0 and 2.0 m/s and dispersed bubble flow at Usl = 5.0 m/s (all for Usg = 0.11 m/s). The results at x/D = 22.1 are also comparable to those of Kocamustafaogullari and Huang (1994) at x/D = 25, with bubble flow occurring in both cases for similar flow conditions.  100 EB/S EB/S  Mg (%)  EB/S  10  B  EB/S  DB  S  EB  EB  B/EB  EB  B DB  5  B/EB  DB  DB  DB  DB  10  EB  Usl(m/s)  EB  0.5 1.0 2.0 3.0 4.0 5.0  B/EB B  B B  DB  4 0  S  S  B  DB DB  EB/S B/EB  B  B B  S  S  15  DB  DB 20  25  x/D  Figure 6.6: Gas mixing index as a function of dimensionless distance downstream for air-water flow at Usg = 0.11 m/s and Dr = 0.208 for various superficial liquid velocities. Letters designate flow regimes: DB: dispersed bubble; B: bubble; EB elongated bubble; S: stratified.  153  6.3.2.2 Vertical gas holdup profiles The gas mixing index indicates the overall uniformity of bubbles, but does not indicate how bubbles are distributed in a cross-sectional plane. Vertical gas holdup profiles were therefore determined to show the cross-sectional distribution. Figure 6.7 illustrates the radial gas holdup profiles along the vertical diameter for each cross-sectional plane along the pipe for Usl = 5.0 m/s and Usg = 0.11 m/s. At x/D = 2.41 (P2) just downstream of gas injection, gas holdup reached a maximum near the bottom of the pipe. The dotted line shows the edge effect caused by the adaptor when the gas phase prevented electrodes from injecting current. At this plane, the gas holdup is expected to be higher, or at least not lower, toward the bottom of the pipe, since it is not far downstream of the injection at the bottom. The peak of gas holdup profiles migrates from near the bottom to the centre of the pipe as x/D increases, leading to less non-uniformity in the cross-section. Figure 6.8 compares the radial gas holdup profiles for various superficial gas velocities at Usl = 5.0 m/s and x/D = 22.1 (P8). The gas holdup increases considerably with increasing superficial gas velocity, especially in the upper part of the pipe. The dotted lines again illustrate the edge effect from the image reconstruction.  154  0.04  x/D 2.41 5.69 8.97 12.2 15.5 18.8 22.1  Local gas holdup  0.03  0.02  0.01  0.00 0  20  40  60  80  Distance from top of pipe (mm)  Figure 6.7: Vertical gas holdup profiles as a function of dimensionless distance downstream for the dispersed bubble air-water flow at Usl = 5.0 m/s, Usg = 0.11 m/s and Dr = 0.208.  0.16  Usg(m/s) 0.11 0.22 0.33 0.44  Local gas holdup  0.12  0.08  0.04  0.00 0  20  40  60  80  Distance from top of pipe (mm)  Figure 6.8: Vertical gas holdup profiles as a function of superficial gas velocity in air-water flow for Usl = 5.0 m/s, Dr = 0.208 and x/D = 22.1 (P8).  155  6.3.2.3 Scale of segregation as a measure of relative size of gaseous entities As noted in section 3.3.2, the scale of segregation, Ls, from equation 3.14 provides a measure of the relative size of gaseous entities in the liquid phase. Figure 6.9 plots correlation coefficients for dispersed bubble air-water flow at Usl = 5.0 m/s, x/D = 22.1 and various superficial gas velocities. Correlation is higher in a cross-sectional plane as Usg increases, likely due to more bubbles enhancing bubble coalescence. The area under the correlograms gave higher Ls with increasing superficial gas velocity, except for Usg = 0.44 m/s where R(r) is similar to that for Usg = 0.33 m/s.  1.0  Usg(m/s) 0.11 0.22 0.33 0.44  0.8  R(r)  0.6  0.4  0.2  0.0 0  2  4  r  6  8  Figure 6.9: Correlation coefficient, R(r), as a function of separation distance, r, for the dispersed bubble air-water flow for Usl = 5.0 m/s, Dr = 0.208, x/D = 22.1 and various superficial gas velocities.  The influence of superficial gas velocity on Ls at Usl = 5.0 m/s is illustrated in Figure 6.10. At Usg = 0.11 m/s, bubbles migrated from the bottom where they were injected to the core of the pipe, and dispersed throughout the cross-sectional plane,  156  leading to smaller bubbles downstream. At higher gas velocities, e.g. 0.22 m/s, more bubbles were present, coalescing and rising to the upper portion the pipe, leading to significantly higher Ls downstream of P5. Ls then became substantially greater at x/D = 22.1 (P8) when the bubbles congregated at the top of the pipe. At the same downstream position (x/D = 22.1), increasing Usg further caused Ls to increase slightly, and Ls was similar for Usg = 0.33 and 0.44 m/s. High-speed camera images at x/D = 22.1 for the same flow conditions confirmed that bubbles were significantly larger for Usg = 0.22 m/s than for Usg = 0.11 m/s. At Usg = 0.33 m/s, the bubbles were more crowded, forming a large void at the top of the pipe. At Usg = 0.44 m/s, more bubbles occupied the cross-section, but the size of gasdominant cluster in the upper part was similar to that for Usg = 0.33 m/s, leading to similar LS.  18  Ls (mm)  16  14  P2 P3  Usg(m/s)  12  0.11 0.22 0.33 0.44  10 0  5  P4  P5  10  P6  15  P8  P7  20  25  x/D  Figure 6.10: Scale of segregation, Ls, as a function of dimensionless distance downstream for dispersed bubble air-water flow at Usl = 5.0 m/s and Dr = 0.208 for various superficial gas velocities.  157  6.3.3 Effect of injection tube geometry Figure 6.11 illustrates the influence of length of the injection tube on mixing quality for the flow of water and a jet-to-pipe diameter ratio of 0.208 at various superficial liquid velocities. The entrance length was approximately 21Dj for the highest velocities examined. The longer injection tube length was sufficient, with injection tube length: diameter ratio, Lj/Dj, of 29, to provide fully developed flow, whereas the flow was not fully developed for the short tube (Lj/Dj = 11). Unlike the liquid jet mixing in water flow, in which mixing was significantly better for a longer tube providing fully-developed jet flow as described in section 4.3.1.2, the gas mixing in air-water flow was almost independent of the injection tube length.  100 EB/S EB/S EB/S  S  S EB/S  EB/S  Mg (%)  B/EB B  B B 10  B  0  EB  B/EB  EB  B B/EB  Usl(m/s) Tube  EB  0.5 0.5 1.0 1.0 2.0 2.0 3.0 3.0 4.0 4.0 5.0 5.0  EB B/EB B  DB DB  5  S  S  EB  B  DB  DB 4  S  B  DB  DB  DB  DB  10  DB 15  B  DB  long short long short long short long short long short long short  DB 20  25  x/D  Figure 6.11: Comparison of effect of long vs. short injection tubes on gas mixing index for air-water flow for Usg = 0.11 m/s, Dj = 0.208 and various superficial liquid velocities. Letters designate flow regimes as in Figure 6.6.  158  The mixing quality was also almost independent of the injection tube diameter when the superficial gas velocity was constant as shown in Figure 6.12. The data consistently showed this behaviour for various superficial liquid velocities and flow regimes. The jet-to-pipe diameter ratio of 0.208 was therefore used for all subsequent tests.  100 EB/S EB/S EB/S  Mg (%) 10  B  0  EB/S  S  DB  EB  B/EB  EB  B DB  DB  5  B  DB  DB  DB  DB 15  B/EB  Usl(m/s) Dr 0.5 0.5 0.5 1.0 1.0 1.0 2.0 2.0 2.0 3.0 3.0 3.0 4.0 4.0 4.0 5.0 5.0 5.0  EB EB B/EB B  B  B  DB  10  S  S  EB  B  DB 4  EB/S  B/EB  B B  S  S  DB 20  DB  0.208 0.104 0.05 0.208 0.104 0.05 0.208 0.104 0.05 0.208 0.104 0.05 0.208 0.104 0.05 0.208 0.104 0.05  25  x/D  Figure 6.12: Comparison of effect of injection tube diameter on gas mixing index for air-water flow at Usg = 0.11 m/s for various superficial liquid velocities. Letters designate flow regimes as in Figure 6.6.  6.3.4 Effects of salt and surfactant on bubble coalescence Coalescence of bubbles occurs as the liquid film separating colliding bubbles drains until it reaches a critical thickness and ruptures. Salts can inhibit bubble coalescence by increasing the surface tension of the film, immobilizing the gasliquid interface and retarding the thinning of this liquid film (Prince and Blanch,  159  1990b). However, the critical NaCl concentration required to immobilize the gasliquid interface of coalescing bubbles in previous work (Lessard and Zieminski, 1971; Prince and Blanch, 1990a; Craig et al., 1993) is 0.078 - 0.175 M, significantly higher than the NaCl concentration in our study (~0.0017 M). The influence of the electrolyte on bubble coalescence in this study is therefore considered to be negligible. Pulp from the pulping process is washed to remove residual chemicals before de-watering and drying to produce the pulp sheets which were the source of the pulp fibres in our study. As a result of the washing, naturally occurring and added surfactants present in industrial pulp suspensions, which are likely to affect bubble coalescence (e.g. Genc et al., 2012), were absent in our study. On the other hand, the tap water used to prepare the suspensions will have contained small amounts of surfactant, as in the great majority of studies of this nature. The differences in surfactant composition and concentrations may have had a small effect on the bubble coalescence behaviour in this study.  6.3.5 Overall gas holdup The overall gas holdup in air-water flow, averaged over seven sensor planes downstream of the injection point, is shown in Figure 6.13. This holdup increased with increasing superficial gas velocity and with decreasing superficial liquid velocity. The gas holdup was slightly higher when the liquid velocity decreased from 5.0 to 4.0 m/s, since the flow was still in the dispersed bubble or bubble flow pattern. The gas holdup increased significantly when the flow pattern changed from  160  bubble flow (Usl = 4.0 m/s) to elongated bubble flow (Usl = 2.0 m/s), and was substantially higher for stratified flow (Usl = 0.5 m/s).  Average gas holdup  0.4  0.3  Usl (m/s)  0.2  0.5 1.0 2.0 3.0 4.0 5.0  0.1  0.0 0.0  0.1  0.2  0.3  0.4  0.5  USg (m/s)  Figure 6.13: Average gas holdup as a function superficial gas and liquid velocities in air-water flow.  Figure 6.14 plots the overall gas holdup in air-suspension flow for various superficial gas velocities and fibre mass concentrations for bubble and dispersed bubble flow patterns at a superficial liquid/pulp velocity of 5.0 m/s. From the suspension flow regime discussed in section 4.3.2, the flow was turbulent for Cm ≤ 1.0% and essentially plug for Cm ≥ 2.0% at Usl = 5.0 m/s. For Cm ≤ 1.0%, the shear stress exceeded the suspension yield stress throughout the suspension, leading to relative motion among fibres and hence turbulent flow, with fibre networks in dynamic equilibrium between simultaneous floc formation and breakdown of flocs by shear (Robertson and Mason, 1957). At Cm = 0.5%, the overall gas holdup was  161  similar to that for water, whereas it was slightly higher at Cm = 1.0%. The denser fibre networks in the turbulent flow regime likely held more bubbles at the higher mass concentration, but the differences in gas holdup were relatively small. The overall gas holdup was significantly higher for plug flow at Cm = 2.0%, since the bubbles were pulled along with the plug. At Cm = 3.0%, the fibre networks in the plug were robust so that the bubbles advanced at nearly the same velocity as the pulp suspension, as shown by the measured gas holdup being similar to the volumetric flow ratio for all superficial gas velocities investigated.  0.10  0.10  Average gas holdup  0.06  0 0.5 1.0 2.0 3.0  0.08  0.06  0.04  0.04  0.02  0.02  0.00 0.0  0.1  0.2  0.3  0.4  Volumetric flow ratio, Gas/total  Cm(%) 0.08  0.00 0.5  Usg (m/s)  Figure 6.14: Average gas holdup for bubble and dispersed bubble flow patterns at Usl = 5.0 m/s for various superficial gas velocities and fibre mass concentrations and comparison with the volumetric flow ratio.  162  6.3.6 Effect of fibre mass concentration on gas mixing in pulp suspension flow Figure 6.15 illustrates the influence of fibre mass concentration on gas mixing in pulp suspension flow at Usl = 5.0 m/s and Usg = 0.11 m/s. At Cm = 0.5%, the bubbles dispersed throughout the pipe cross-section in turbulent suspension flow, resulting in mixing quality similar to that for water for dispersed bubble flow. At Cm = 1.0%, the flow was turbulent, with more dense fibre networks forming in dynamic equilibrium with disintegration of flocs by shear, accelerating bubble coalescence downstream of P5. Larger bubbles then occupied the top of the pipe, as shown by tomographic images in Figure 6.16b, leading to significantly worse mixing downstream. The mixing quality deteriorated substantially at higher Cm for plug flow, likely due to strong fibre networks, preventing bubbles from migrating to the centre of the pipe and leading to non-uniform distribution of bubbles in the cross-section. At Cm = 2.0%, the bubbles moved further from the pipe wall downstream of the injection point, resulting in slightly better mixing downstream from P2 to P4, and appeared to reach the top by skirting around the centre of the pipe as illustrated in Figure 6.16c, with strong fibre networks and gravity promoting bubble coalescence, leading to considerably worse mixing downstream of P5. At Cm = 3.0%, the mixing quality did not vary significantly along the pipe, and was consistently worse than for lower mass concentrations at each downstream position. Due to robust fibre networks in the core of the pipe, the bubbles travelled as a cluster close to the pipe wall, sliding around the pipe periphery, from bottom to  163  top of the pipe, as illustrated in Figure 6.16d, resulting in non-uniform gas dispersion and poor mixing.  100  Cm(%)  Mg (%)  0 0.5 1.0 2.0 3.0  10 P2 P3  P4  P5  P6  P7  P8  3 0  5  10  15  20  25  x/D  Figure 6.15: Gas mixing index as a function of dimensionless distance downstream for Usl = 5.0 m/s, Usg = 0.11 m/s and various fibre mass concentrations. Figure 6.17 plots vertical gas holdup profiles along a diameter at x/D = 12.2 (P5) and 22.1 (P8) for the same flow conditions, with different fibre mass concentrations. At x/D = 12.2, the profiles for water and dilute (Cm = 0.5%) fibre suspension reach peaks near the centre of the pipe, whereas the peak shifts slightly toward the top of the pipe for Cm = 1.0%, showing that the bubbles began to coalesce and rise to the upper portion, as illustrated in Figure 6.17a. The vertical profile for Cm = 2.0% shows more asymmetry, likely due to strong fibre networks in the core of the pipe. The profile for Cm = 3.0% differs significantly from those for lower Cm. The peak of gas holdup profile is at the bottom of the pipe in this case,  164  Top  Figure 6.16: Tomographic images for air-pulp suspension flow at Usl = 5.0 m/s and Usg = 0.11 m/s for: (a) Cm = 0.5% (b) Cm = 1.0% (c) Cm = 2.0% (d) Cm = 3.0%. Locations of planes P2, P4, P6 and P8 are shown in Figure 3.2. Tomographic images for air-water flow for the same flow conditions are illustrated in Figure 6.3a.  indicating that robust fibre networks in the core prevented bubbles from passing through the centre, causing them to cluster at the bottom of the pipe. At x/D = 22.1,  165  the profiles for water and 0.5% pulp suspension are similar to those at x/D = 12.2, with the peaks shifting slightly toward the top of the pipe, as shown in Figure 6.17b, indicating  developing  gas-liquid/suspension  flow.  At  higher  fibre  mass  concentrations, the peak of gas holdup profiles approach the top of the pipe, showing the influence of the fibre networks on accelerating bubble coalescence, with substantially higher peak for Cm ≥ 2.0% for plug flow. Figure 6.18 compares the scale of segregation for Usl = 5.0 m/s, Usg = 0.11 m/s and various mass concentrations. Ls was similar for the first two planes since the bubbles concentrated at the bottom of the pipe downstream of the injection point. Ls was then lower for Cm = 0 and 1.0% since small bubbles dispersed throughout the cross-section, whereas Ls for Cm = 3.0% did not vary significantly since the bubbles travelled as a cluster adjacent to the pipe wall. Ls for Cm = 1.0% was considerably greater downstream of P5, likely due to bubbles coalescing and migrating to the top of the pipe, resulting in a similar average bubble size to that for Cm = 3.0%. Figure 6.19 illustrates the influence of fibre mass concentration on mixing at a lower Usl, 2.0 m/s. Mixing differed significantly for the first two planes downstream of the injection, but was similar downstream of P4 where elongated bubble flow was approached. The gas, introduced to the main flow just upstream of P2, penetrated less for higher mass concentrations, likely due to the hindering resistance caused by the strong fibre networks in the core of the pipe. This is also shown by the tomographic images at P2 in Figure 6.20. As a result, the gas mixing index was substantially higher for higher concentrations. Bubbles then migrated to the centre  166  (a)  Local gas holdup  0.05  0.04  Cm(%)  0.03  0 0.5 1.0 2.0 3.0  0.02  0.01  0.00 0  20  40  60  80  Distance from top of pipe (mm)  (b) 0.08  Cm(%) 0 0.5 1.0 2.0 3.0  Local gas holdup  0.06  0.04  0.02  0.00 0  20  40  60  80  Distance from top of pipe (mm)  Figure 6.17: Vertical gas holdup profiles as a function of fibre mass concentration for Usl = 5.0 m/s, Usg = 0.11 m/s at: (a) x/D = 12.2 and (b) x/D = 22.1.  167  18  Ls (mm)  16  14  P2  P3  Cm(%) 12  P4  P5  0 1.0 3.0  P6  P8  P7  10 0  5  10  15  20  25  x/D  Figure 6.18: Scale of segregation as a function of dimensionless distance downstream for Usl = 5.0 m/s, Usg = 0.11 m/s and various fibre mass concentrations.  of the pipe for Cm ≤ 2.0%, whereas the robust fibre networks for Cm = 3.0% caused the bubbles to concentrate near the bottom of the pipe, as illustrated by the high gas-concentration regions at P3 in Figure 6.20d, leading to profoundly higher gas mixing index for Cm = 3.0%. At this superficial liquid/pulp velocity, the buoyancy was sufficient that the bubbles rose through the centre of the pipe for Cm = 2.0%, whereas they skirted around the core of the pipe at Usl = 5.0 m/s for the same concentration (see Figure 6.16c). The robust fibre networks for Cm ≥ 2.0% also retarded the arrival of bubbles at the top of the pipe, as shown by the tomographic images at P4. Bubbles continued to coalesce along the pipe, finally becoming elongated, with similar mixing quality for different Cm downstream of P4.  168  100 90 80 70 60  Cm(%) 0 0.5 1.0 2.0 3.0  50  Mg (%)  40  B/EB  B  30  B  B 20  B P2  10 0  DB  EB  P7  P8  B P5  B  B/EB EB  P6  P4  P3 5  10  15  20  25  x/D  Figure 6.19: Gas mixing index as a function of dimensionless distance downstream for Usl = 2.0 m/s, Usg = 0.11 m/s and various fibre mass concentrations. Letters designate flow regimes as indicated in Figure 6.6.  Figure 6.21 plots the scale of segregation for the same flow conditions and various fibre mass concentrations. Ls at P2 for Cm ≤ 0.5% was lower than for higher Cm, i.e. the bubbles were smaller, likely due to greater gas penetration, whereas the bubbles spread less and formed a gas cluster at the bottom of the pipe for higher Cm. Ls was then similar at P3 for Cm ≤ 2.0%, since the bubbles migrated to the core of the pipe and dispersed in the cross-section, whereas the robust fibre networks in the core of the pipe for Cm = 3.0% caused the bubbles to cluster near the pipe wall, leading to significantly higher Ls. The bubbles then rose to the upper part of the pipe and coalesced for Cm ≤ 1.0%, resulting in significantly higher Ls at P4 than for the two proceeding planes. The bubbles, however, did not reach the top of the pipe at  169  Top  Figure 6.20: Tomographic images for air-pulp suspension flow at Usl = 2.0 m/s, Usg = 0.11 m/s for: (a) Cm = 0% (b) Cm = 1.0% (c) Cm = 2.0% (d) Cm = 3.0%. Locations of planes P2 – P5 are shown in Figure 3.2. Colour scales indicate gas holdups  P4 for Cm ≥ 2.0%, likely due to the dense fibre networks in the core of the pipe, and Ls was slightly lower than for lower Cm. Ls was then similar downstream of P4 when the elongated bubble flow was approached.  170  18  Ls (mm)  16  14 P4  P5  P6  P7  P8  Cm(%)  P2 P3  0 0.5 1.0 2.0 3.0  12  10 0  5  10  15  20  25  x/D  Figure 6.21: Scale of segregation, Ls, as a function of dimensionless distance downstream for Usl = 2.0 m/s, Usg = 0.11 m/s and various fibre mass concentrations.  Figure 6.22 plots the gas mixing index at Usl = 0.5 m/s and Usg = 0.11 m/s for various fibre mass concentrations. The trend is opposite to that for Usl = 5.0 m/s and the same Usg (compare Figure 6.15), i.e. mixing for Usl = 0.5 m/s improved with increasing Cm. The flow pattern was stratified flow for Cm ≤ 0.5%, with poor mixing downstream, since buoyancy was dominant. At Cm = 1.0%, the stronger fibre networks confined the gas to the upper part of the pipe and caused liquid/pulp slugs to advance at the top of the pipe, leading to elongated bubbles, with significantly better mixing than for lower mass concentrations. The buoyancy was more significant downstream of P4, so that the slugs were less frequent, stratified flow was approached, and mixing was considerably worse. For higher mass  171  concentrations (Cm ≥ 2.0%), the robust fibre networks created more frequent liquid/pulp slugs, with substantially better mixing than for lower Cm.  100 80  S EB/S EB/S  60  EB/S EB EB/S  Mg (%)  40  P2  EB P3  S  S  S  S  EB  EB  EB  EB  EB  EB  P5  P6  P7  EB  EB EB P4  EB P8  Cm(%) 0 0.5 1.0 2.0 3.0  20  10 0  5  10  15  20  25  x/D  Figure 6.22: Gas mixing index as a function of dimensionless distance downstream for Usl = 0.5 m/s, Usg = 0.11 m/s and various fibre mass concentrations. Letters designate flow regimes as indicated in Figure 6.6.  6.3.7 Flow regime for air-suspension flow and effect of superficial liquid/pulp velocity Flow regimes for air-suspension horizontal flow for x/D = 22.1 and various fibre mass concentrations are summarized in Table 6.3. For dilute suspensions (Cm ≤ 0.5%) at Usg = 0.11 m/s, the flow regime varied from stratified to elongated bubble, bubble and dispersed bubble flow as the superficial liquid/pulp velocity increased. For Cm ≥ 1.0%, the flow regime did not vary much: only elongated bubble and bubble flow existed. At a higher superficial gas velocity, 0.44 m/s, the flow  172  regime was similar to that for Usg = 0.11 m/s, with some differences likely due to the presence of more bubbles. For Usl = 5.0 m/s and Cm ≤ 0.5%, bubbles congregated at the top of the pipe for Usg = 0.44 m/s, whereas they were dispersed more uniformly for Usg = 0.11 m/s. For Usl = 1.0 m/s and Cm ≤ 0.5%, stratified flow prevailed for Usg = 0.44 m/s, whereas elongated bubble flow was found for Usg = 0.11 m/s. For higher fibre mass concentrations (Cm ≥ 2.0%), transition from bubble to elongated bubble flow occurred at Usl = 4.0 m/s for Usg = 0.44 m/s, whereas this transition occurred at a lower liquid/pulp velocity (Usl = 3.0 m/s) for Usg = 0.11 m/s.  Table 6.3: Flow regimes (identified as in Figure 6.6) for air-suspension horizontal flow for Usg = 0.11 and 0.44 m/s, x/D = 22.1 and various fibre mass concentrations  Cm (%)  0.5  1.0  0 0.5 1.0 2.0 3.0  S S EB EB EB  EB EB EB EB EB  0 0.5 1.0 2.0 3.0  S S EB EB EB  S S EB EB EB  Usl (m/s) 2.0 3.0 Usg = 0.11 m/s EB B/EB EB B/EB EB B/EB EB B/EB EB B/EB Usg = 0.44 m/s EB B/EB EB B/EB EB B/EB EB EB EB EB  4.0  5.0  B B B B B  DB DB B B B  B B B B/EB B/EB  B B B B B  Figure 6.23 illustrates the influence of superficial liquid/pulp velocity on mixing quality for Usg = 0.11 m/s, x/D = 22.1 and various fibre mass concentrations. The gas uniformity improved substantially with increasing Usl for Cm ≤ 0.5% due to  173  different flow regimes as shown in Table 6.3. Mixing for Cm ≥ 2.0%, however, improved slightly as Usl increased, since the flow regime only changed from elongated bubble to bubble flow. At Usl = 0.5 m/s, denser fibre networks for these concentrations caused liquid/pulp slugs at the top of the pipe, resulting in elongated bubble flow, whereas stratified flow occurred at lower mass concentrations (Cm ≤ 0.5%) since buoyancy was the dominant force. For Usl ≥ 4.0 m/s, a rigid plug in the core of the pipe, together with buoyancy, caused bubbles to congregate at the top of the pipe for Cm ≥ 2.0%, whereas the bubbles dispersed more uniformly in the cross-section for Cm ≤ 0.5%. Superficial liquid/pulp velocity thus did not have a strong influence on mixing for Cm ≥ 2.0%. At Cm = 1.0%, mixing depended strongly on Usl, even though the flow regime only varied from elongated bubble to bubble flow, since the slug frequency was relatively low, with stratified flow approached. Mixing therefore deteriorated substantially with decreasing Usl.  80  Cm(%) 0 0.5 1.0 2.0 3.0  Mg (%)  60  40  20  0 0  2  4  6  Usl (m/s)  Figure 6.23. Gas mixing index as a function of superficial liquid velocity for Usg = 0.11 m/s, x/D = 22.1 (plane P8) and various fibre mass concentrations.  174  Figure 6.24 illustrates the flow regimes and corresponding mixing quality for each downstream position for Usg = 0.11 m/s. The flow regime for each downstream position for Cm = 0.5% was similar to that for water, with the flow regime and mixing quality varying greatly with position downstream and Usl, as shown in Figure 6.6. There was, however, less variation in flow regime and mixing quality for Cm = 1.0% as illustrated in Figure 6.24a. For Usl ≥ 2.0 m/s, mixing was poor at P2, as bubbles concentrated at the bottom of the pipe near the injection point. Mixing then improved significantly at P3 as bubbles spread throughout the cross-section for the dispersed bubble (DB) flow. The bubbles reached the top of the pipe and coalesced faster downstream for lower Usl, with considerably worse mixing for bubble (B) and elongated bubble (EB) flow. For Usl ≤ 1.0 m/s, bubbles migrated rapidly to the top of the pipe after their injection, forming elongated bubbles, with lower slug frequency and worse mixing than for Usl = 2.0 m/s. Figure 6.24b shows the flow regimes and corresponding mixing quality for Cm = 3.0%. Mixing quality was clearly less dependent on Usl for each downstream position, likely due to the robust fibre networks in the core of the pipe causing bubbles to concentrate near the wall, even at high Usl. For Usl ≥ 4.0 m/s, bubbles slid around the pipe periphery without passing through the core of the pipe (e.g. see Figure 6.16d for Usl = 5.0 m/s and Cm = 3.0%), resulting in consistently worse mixing quality along the pipe. At lower Usl, buoyancy caused more bubbles to concentrate in the upper part of the pipe, finally forming elongated bubbles for Usl ≤ 2.0 m/s, with relatively high liquid/pulp slug frequency due to robust fibre networks. Mixing therefore did not vary significantly with Usl and downstream position.  175  (a) 100 EB/S  EB EB/S  EB EB/S  Mg (%)  B  10  B  DB  EB  EB/S  EB  EB  EB EB EB  Usl(m/s)  B/EB  0.5 1.0 2.0 3.0 4.0 5.0  EB B/EB  B  B  DB DB DB  EB  B/EB EB B  B  B  EB  B  DB  B  DB DB  DB DB  DB  B  B  4 0  5  10  15  20  25  x/D  (b) 100 80 60 EB/S EB EB B  Mg (%)  40 B  B 20  10  0  5  EB  EB  EB  B B  EB B B  B  B  10  EB B B B  15  EB  EB  EB B/EB  EB  B B  20  Usl(m/s)  B/EB B B  0.5 1.0 2.0 3.0 4.0 5.0  25  x/D  Figure 6.24: Gas mixing index as a function of dimensionless distance downstream for Usg = 0.11 m/s, Dr = 0.208 and various superficial liquid velocities at: (a) Cm = 1.0% and (b) Cm = 3.0%. Letters designate flow regimes as indicated in Figure 6.6.  176  6.3.8 Effect of superficial gas velocity on mixing in pulp suspensions Figure 6.25 compares the gas mixing index at Usl = 5.0 m/s and x/D = 22.1 for various fibre mass concentrations and superficial gas velocities. Mixing depended strongly on Usg for all mass concentrations investigated. Uniformity of gas deteriorated with increasing Usg, likely due to more bubbles coalescing and occupying the top of the pipe. Mixing in pulp suspension flow at Cm = 0.5% was similar to that for water for all superficial gas velocities examined, since the suspension flow was turbulent at this mainstream velocity. The slightly worse mixing than for water, especially for higher Usg, was likely due to loose fibre networks causing bubbles to segregate to the top of the pipe. Mixing was worse for Cm = 1.0%, likely because denser fibre networks promoted bubble coalescence, and deteriorated considerably for plug flow with Cm ≥ 2.0%. Figure 6.26 illustrates the influence of superficial gas velocity on mixing for Usl = 2.0 m/s, x/D = 22.1 and various fibre mass concentrations. The flow regime was elongated bubble flow for all conditions tested. The mixing quality decreased substantially as the superficial gas velocity increased, due to larger bubbles and lower frequency of liquid/pulp slugs, consistent for various Cm.  177  50  Cm(%) 0 0.5 1.0 2.0 3.0  40  Mg (%)  30  20  10  0 0.0  0.1  0.2  0.3  0.4  0.5  Usg (m/s)  Figure 6.25: Gas mixing index as a function of superficial gas velocity for Usl = 5.0 m/s, x/D = 22.1 (plane P8) and various fibre mass concentrations.  60  Cm(%)  Mg (%)  50  40  0 0.5 1.0 2.0 3.0  EB  EB EB  EB 30  20 0.0  0.1  0.2  0.3  0.4  0.5  Usg (m/s)  Figure 6.26: Gas mixing index as a function of superficial gas velocity for Usl = 2.0 m/s, x/D = 22.1 (plane P8) and various fibre mass concentrations. EB indicates the elongated bubble flow regime.  178  Figure 6.27 shows the effect of superficial gas velocity on mixing quality at a lower Usl, 0.5 m/s, for various fibre mass concentrations. Mixing was almost independent of Usg for stratified flow with Cm ≤ 0.5%. Mixing, however, deteriorated with increasing Usg for higher mass concentrations corresponding to elongated bubble flow. At Cm = 1.0%, liquid/pulp slugs were less frequent, and stratified flow was approached, leading to profoundly worse mixing than for Cm ≥ 2.0% for all superficial gas velocities investigated.  80 S  S  S  70  EB  EB  60  Mg (%)  S  EB  EB  Cm(%) EB  50 40  0 0.5 1.0 2.0 3.0  EB  EB  EB EB  30 20 0.0  0.1  0.2  0.3  0.4  0.5  Usg (m/s)  Figure 6.27: Gas mixing index as a function of superficial gas velocity for Usg = 0.5 m/s, x/D = 22.1 (plane P8) and various fibre mass concentrations. Letters identify flow regimes as indicated in Figure 6.6.  Supplementary data relevant to this chapter are provided in Appendix C.  179  6.4 Conclusions Mixing of gas into water and pulp suspension flow in a horizontal pipe was investigated for a range of operating conditions based on electrical resistance tomography (ERT). ERT measurement with resistor adaptors resolved anomalous voltage data and reconstructed images where electrodes lost electrical contact when the gas phase occupied the top of the pipe, providing a by-pass route for the electrical current. The gas mixing index indicated worse mixing in air-water flow at Usl = 5.0 m/s when bubbles concentrated at the bottom of the pipe downstream of injection, when they coalesced and migrated to the top of the pipe as they were carried downstream, and with increasing superficial gas velocity due to more and larger bubbles. Mixing quality was considerably lower when the superficial liquid velocity decreased as the flow pattern changed from dispersed bubble to bubble, elongated bubble and stratified flow. Vertical gas holdup profiles reached a maximum at the centre of the pipe for dispersed bubble flow, and the peaks shifted toward the top of the pipe, with higher peaks at higher superficial gas velocities. The scale of segregation, Ls, was considerably higher when bubbles coalesced and occupied the top of the pipe than for dispersed bubble flow, but did not vary significantly when the bubbles clustered at the top of the pipe. Gas dispersion in horizontal pulp suspension flow depends on the flow regime and fibre mass concentration in a complex manner. Gas dispersion in pulp suspension flow for the bubble flow regime at Usl = 5.0 m/s depended substantially  180  on fibre mass concentration. For Cm ≤ 1.0%, the uniformity of bubbles was similar to that for water since the flow was turbulent, but was profoundly worse for higher concentrations (Cm ≥ 2.0%) when the flow was plug, as robust fibre networks in the core caused bubbles to concentrate near the pipe wall. The fibre mass concentration, however, did not have a profound influence on mixing when the elongated bubble flow was approached at Usl = 2.0 m/s. At Usl = 0.5 m/s, the mixing quality was better with increasing Cm as the flow regime changed from stratified flow for Cm ≤ 0.5% to elongated bubble flow for Cm ≥ 1.0%, since robust fibre networks caused liquid/pulp slugs at the top of the pipe. The superficial liquid/pulp velocity affected gas uniformity profoundly for Cm ≤ 0.5% since it caused the flow regime to vary significantly, whereas it had less influence for Cm ≥ 2.0% because only two flow regimes, bubble and elongated bubble flow, existed for different Usl. For Cm = 1.0%, mixing deteriorated considerably with decreasing Usl, due to less frequent liquid/pulp slugs for elongated bubble flow, approaching stratified flow. Gas uniformity in pulp suspension flow decreased with increasing superficial gas velocity for dispersed bubble and bubble flow, since more bubbles congregated at the top of the pipe. For elongated bubble flow, mixing deteriorated as Usg increased because of larger bubbles and lower liquid/pulp slug frequency. Mixing was almost independent of Usg for stratified flow.  181  7. Gas Dispersion in Pulp-Suspension Flow in the Presence of an In-Line Mechanical Mixer8  7.1 Introduction Gas bleaching processes are usually carried out in the medium-consistency range, with fibre mass concentration from 8 to 16%, and gas void fraction between 0.11 and 0.26 (Bennington, 1996; Bennington, 2004). Mixing is usually provided by high-shear mixers ahead of tower reactors. Gaseous chemicals are injected and mixed with pulp suspensions in zones of intense shear created by high rotational speed across narrow gaps, causing turbulence and uniform gas dispersion. The gas-suspension mixture then flows along the pipes and enters the tower reactors. Few studies have been carried out to investigate mixing of gas into pulp suspensions. Bennington (1993) evaluated gas dispersion in a batch-operated shear tester based on the bulk flow pattern and power dissipation. Accumulation of gas around the rotor was observed when sufficient gas was introduced, leading to a sudden drop in the required power. This is referred to as “flooding”. When this occurs, the rotor is unable to transfer momentum to the suspension, causing the suspension flow to stop, although the rotor continues to rotate in the gas pocket.  8  A version of this chapter has been submitted for publication: Yenjaichon, W., Grace, J.R., Lim, C.J.,  Bennington, C.P.J., 2012. Gas dispersion in pulp-suspension flow in the presence of an in-line mechanical mixer.  182  This poor mixing occurred at lower void fractions as the suspension mass concentration increased, since higher yield stress prevented contact between the suspension and rotor, promoting gas segregation. Smith and Bennington (1995) extended this work to a continuous shear mixer, with similar results, and suggested that the gas volume fraction, calculated based on the ratio of volumetric flow of gas to total flow, should be kept below 10% to ensure effective mixing for 3%wt. pulp suspension. Kourunen et al. (2011) evaluated gas mixing in pulp suspension with an MC mixer based on the conductivity distribution in the cross-section, but few trials were performed. More experimental data are needed to provide better understanding of gas mixing in pulp suspensions. Flow regimes for gas-liquid horizontal flow have been identified by several authors, with details provided in section 6.1. In this chapter, four flow regimes stratified flow (S), elongated bubble flow (EB), bubble flow (B) and dispersed bubble flow (DB) - were observed. The aim is to investigate the distribution of local gas holdup in cross-sectional planes downstream of an in-line mechanical mixer for various fibre mass concentrations, superficial gas and liquid/pulp velocities, and impeller speeds, and hence to establish optimum operating conditions.  7.2 Experimental Details Details of the experimental flow loop, test section and in-line mechanical mixer system are presented in section 3.1. The tests were conducted using the procedure provided in section 3.2.2, with the degree of mixing quantified by the gas  183  mixing index, defined in section 3.3.2. The fibre mass concentration varied from 0 to 3.0%, superficial liquid/pulp velocities from 0.5 – 3.0 m/s, superficial gas velocities from 0.055 – 0.44 m/s and impeller speeds from 0 – 800 rpm. The injection tube (Dr = 0.208) was sufficiently long to provide fully-developed gas flow, with a length:diameter ratio of 29, whereas the entry length is estimated to be approximately 21Dj for the highest velocity investigated.  7.3 Results and Discussion 7.3.1 Effect of velocities and rotational speed on gas mixing in water flow The mainstream Reynolds number varied from 38,100 to 229,000, and the side-stream Reynolds number from 2,700 to 10,800. Figure 7.1a illustrates the effect of superficial gas velocity on mixing for Usl = 3.0 m/s and N = 400 rpm. Air was injected at the bottom of the pipe ahead of the impeller, just upstream of P2, in the shaded entry region. Mixing in this region is complex, given that both gas injection and impeller are present. At Usg = 0.11 m/s, the impeller dispersed the bubbles uniformly in the cross-section, and the mixing quality improved from P2 to P4, as shown by a decrease in gas mixing index. Bubble coalescence and migration then became dominant, resulting in worse mixing downstream, with concentrated bubbles at the top at P8 for bubble flow. The flow pattern was similar for each  184  superficial gas velocity tested (0.11 – 0.44 m/s), with more rapid bubble coalescence at the top of the pipe and worse mixing as Usg increased. Figure 7.1b shows the influence of Usg on mixing at a lower Usl, 1.0 m/s. Mixing was almost independent of Usg for Usg ≥ 0.22 m/s as stratified flow was approached. There was, however, a great difference in mixing quality for Usg = 0.11 and 0.22 m/s, indicating that there is a gas volume fraction (ratio of gas to total volumetric flow rate) that the system can handle between these two velocities. The impeller dispersed gas more uniformly in the cross-sectional P2 plane for Usg = 0.11 m/s. Introducing more gas into the system caused a gas cavity to form around the impeller, significantly increasing gas holdup and causing deterioration in mixing quality. This sudden change occurred at a gas volume fraction of 0.18, very similar to the gas fraction of 0.2 observed for flooding found during air-water mixing in a continuous laboratory mixer (Smith and Bennington, 1995). Figure 7.2 plots the gas mixing index for Usg = 0.11 m/s, N = 400 rpm and various superficial liquid velocities. Bubbles dispersed uniformly throughout the P2 plane for Usl ≥ 1.0 m/s, whereas buoyancy caused the gas to segregate rapidly to the top of the pipe for Usl = 0.5 m/s, leading to stratified flow and poor mixing. The mixing quality for Usl = 1.0 m/s at P2 was significantly lower than for higher Usl, likely due to less turbulent shear on bubbles, resulting in larger bubbles and less uniformity. Mixing deteriorated greatly downstream with decreasing Usl since the flow regime changed from dispersed bubble to bubble and ultimately to stratified flow.  185  (a) 100  Usg (m/s) 0.11 0.22 0.33 0.44 B  Entry Region  Mg (%)  DB DB  DB  DB  DB  DB  10  DB  P2 5  DB  Impeller P3 0  DB DB DB DB P4  5  B DB  B  B B  B  B B  DB  DB DB  B B  DB DB  P8  P7  P6  P5  10  15  20  25  x/D  (b) 100 90 80 70 60  EB/S  EB/S  50  EB/S  40  Mg (%)  S  S  S  EB/S  EB/S  S  S EB/S  EB/S  B  30  Usg (m/s)  DB  0.11 0.22 0.33 0.44  20  Impeller 10 0  5  10  15  20  25  x/D  Figure 7.1: Gas mixing index and flow regimes as a function of dimensionless distance downstream, x/D, in air-water flow for N = 400 rpm and various superficial gas velocities for: (a) Usl = 3.0 m/s; (b) Usl = 1.0 m/s. P followed by a number designates sensor plane shown in Figure 3.5. Letters identify flow regimes as in Figure 6.6.  186  100 S  EB/S S  EB/S  EB/S  EB/S  S  S  S  S  EB/S  EB/S B  B  B  Usl(m/s)  Mg (%)  B DB  DB DB  10  B DB  DB DB  DB  DB  DB  0.5 1.0 2.0 3.0  DB  DB  Impeller 5  0  5  10  15  20  25  x/D  Figure 7.2: Gas mixing index and flow regimes as a function of dimensionless distance downstream in air-water flow for Usg = 0.11 m/s, N = 400 rpm and various superficial liquid velocities. Letters identify flow regimes as in Figure 6.6. Figure 7.3 shows the influence of impeller speed on mixing quality at Usg = 0.11. Two configurations of the impeller when stationary ‒parallel and perpendicular to the flow‒ were also investigated. At Usl = 3.0 m/s, the mixing quality at plane P2 for a rotating impeller was slightly better than for static impellers and for a tee mixer alone, as shown in Figure 7.3a. The bubbles coalesced and migrated rapidly to the top of the pipe downstream of P4 for the tee mixer alone, whereas bubbles dispersed throughout the cross-section and began to congregate at the upper part of the pipe downstream of P7 for the rotating impeller, leading to much better mixing. Mixing was, however, almost independent of impeller speed, likely due to low residence time in the high-shear zone, as discussed in section 5.3.1 (Figure 5.4). Downstream turbulence caused by the parallel static mixer was likely higher  187  than for the tee alone, but less than for both the rotating impeller and the perpendicular stationary impeller. The effect of rotational speed on mixing at a lower Usl, 0.5 m/s, is illustrated in Figure 7.3b. At P2 and P3, mixing with the impeller was worse than without it. For tee mixing alone, bubbles migrated rapidly from the injection point at the bottom of the pipe and reached the top at P2. They then disintegrated into smaller bubbles, occupying the top of the pipe at P3, with stratified flow approached at P4, as described in section 6.3.2.1. In the presence of the impeller, introducing gas into the main stream at this low Usl caused a gas cavity to form around the impeller, occurring at a gas volume fraction of 0.18. Buoyancy also caused gas to segregate rapidly to the top of the pipe, leading to profoundly worse mixing than for the tee alone, with stratified flow approached immediately downstream of P2. Mixing with the impeller was almost independent of rotational speed and mixer configurations, and was similar to that without it when stratified flow was approached. The gas mixing index reached a similar value of ~0.7, whereas a value closer to 1.0 was expected since the gas segregated completely to the top of the pipe, likely due to an image-reconstruction edge effect, as discussed in section 6.3.1.  7.3.2 Effect of fibre mass concentration on gas mixing in pulp suspensions Figure 7.4 plots the gas mixing index for Usl = 3.0 m/s, Usg = 0.11 m/s, N = 600 rpm and various fibre mass concentrations. In the high-shear (P2) zone  188  (a) 100  N (rpm)  Mg (%)  NA 0 (//) 0( ) 210 401 595 803 P2 B  P3 DB  DB  DB DB  Impeller DB 0  P5  P4  10  5  P6  DB  5  B  P8  P7 B/EB  B/EB B  B B DB  DB  DB  10  B  DB DB  15  20  25  x/D  (b) 100 90 80 EB/S S  70  S  S  S  S  S  60  Mg (%)  S  N (rpm)  EB/S  NA 0 (//) 0( ) 205 402 600 801  EB/S  50  40  Impeller 30 0  5  10  15  20  25  x/D  Figure 7.3: Gas mixing index and flow regimes as a function of dimensionless distance downstream in air-water flow for Usg = 0.11 m/s and various impeller speeds at: (a) Usl = 3.0 m/s; (b) Usl = 0.5 m/s. Letters identify flow regimes as in Figure 6.6. // and  indicate parallel and normal to flow.  immediately downstream of the impeller, the energy from the impeller and the flow was sufficient to disperse bubbles uniformly across the pipe, with mixing quality in  189  suspension similar to that in water. Mixing became significantly worse downstream as the fibre mass concentration increased, likely due to decaying turbulence accompanied by reflocculation. For Cm = 0.5%, the suspension flow was turbulent, with mixing quality similar to water for P2 to P4. Mixing was then slightly worse than for water, likely due to the fibre networks in turbulent suspension flow accelerating bubble coalescence. For Cm ≥ 2.0%, reflocculation likely occurred more rapidly, as shown by a higher gas mixing index downstream of P2. The stronger fibre networks in the core of the pipe caused bubbles to migrate to the top, leading to bubble and elongated bubble flow downstream, with substantially worse mixing than at lower Cm, for which only dispersed bubble and bubble flow occurred.  100  Cm(%)  Mg (%)  0 0.5 1.0 2.0 3.0  P2 DB  10  5  P3 DB  DB Impeller DB 0  5  P4 DB DB DB DB  P6 B  P5 B  B  DB DB  DB  P7 B/EB B/EB B DB  B/EB B B B  DB DB  P8  DB  DB  10  15  20  25  x/D Figure 7.4: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Usl = 3.0 m/s, Usg = 0.11 m/s, N = 600 rpm and various fibre mass concentrations. P2 – P8 identify planes, whereas other letters identify flow regimes.  190  Figure 7.5 illustrates the influence of fibre mass concentration on mixing at a lower Usl, 0.5 m/s. Buoyancy was significant at this superficial liquid/pulp velocity, and the dependence of mixing quality on fibre mass concentration was opposite to that at Usl = 3.0 m/s, i.e. mixing improved with increasing Cm, likely due to liquid/pulp slugs caused by robust fibre networks at higher mass concentrations. For Cm ≤ 0.5%, the mixing behaviour was similar to that without an impeller (Figure 6.22 in section 6.3.6), i.e. buoyancy was likely dominant so that the gas segregated to the upper part of the pipe immediately downstream of the impeller, leading to stratified flow and poor mixing. For Cm ≥ 2.0%, the gas also segregated immediately downstream of the impeller, but strong fibre networks caused liquid/pulp slugs to advance along the upper part of the pipe, leading to elongated bubbles and considerably improved mixing. For Cm = 1.0%, the imposed shear was sufficient to disintegrate fibre networks, leading to stratified flow and mixing quality similar to that for Cm ≤ 0.5%, whereas liquid/pulp slugs caused by fibre networks occurred for the same flow conditions without the impeller, as shown in Figure 6.22. For Cm = 2.0%, the energy supplied was sufficient to disrupt the fibre networks in the high-shear (P2) zone, with mixing quality similar to lower concentrations. Reflocculation then likely occurred rapidly since the energy from the impeller was not sustained, causing liquid/pulp slugs to form immediately downstream, resulting in profoundly better mixing. For Cm = 3.0%, fibre networks likely re-established more rapidly downstream of the impeller, leading to significantly better mixing at P2 than for lower Cm. Mixing with the impeller was then almost independent of Cm downstream  191  of P3, where elongated bubble flow was approached for Cm ≥ 2.0%, and was similar to that for a tee mixer alone, indicating rapidly decaying turbulence.  100 80  P3  P2  EB/S S  60  P4  P5  P6  P7  P8  S  S  S  S  S  EB/S 40  Mg (%)  EB/S  EB  EB  EB  EB  EB  EB  Cm(%) 0 0.5 1.0 2.0 3.0 3.0 (w/o impeller)  20  Impeller 10 0  5  10  15  20  25  x/D Figure 7.5: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Usl = 0.5 m/s, Usg = 0.11 m/s, N = 600 rpm and various fibre mass concentrations.  Figure 7.6 shows the effect of fibre mass concentration on mixing for Usl = 1.0 m/s, Usg = 0.11 m/s and N = 600 rpm. At the corresponding gas volume fraction of 0.1, gas cavity formed around the impeller leading to flooding at Cm ≥ 1.0%, resulting in considerably worse mixing than for Cm ≤ 0.5% in the P2 plane. This poor mixing occurred at a lower gas volume fraction for higher Cm, due to higher suspension yield stress preventing contact between suspension and impeller, leading to the gas cavity around the impeller (Smith and Bennington, 1995). For Cm ≤ 0.5%, the impeller dispersed bubbles uniformly in the P2 plane, with bubbles  192  coalescing and migrating to the top of the pipe downstream. Hence mixing worsened as the flow pattern changed from dispersed bubble to bubble and then to a transition between elongated bubble and stratified flow. For Cm ≥ 1.0%, the imposed shear was sufficient to disintegrate the fibre networks, leading to gas segregation to the top of the pipe, i.e. stratified flow. The re-forming of robust fibre networks likely occurred downstream for Cm ≥ 2.0%, causing liquid/pulp slugs at the top of the pipe and hence improved mixing downstream. The fibre networks for Cm = 1.0% were, however, too weak to create slugs, and the flow remained stratified. For Cm = 3.0%, the fibre networks were robust, likely causing more rapid reflocculation than for Cm = 2.0%, as shown by a sudden drop in Mg downstream of P3 for Cm = 3.0%, whereas Mg decreased gradually downstream of P5 for Cm = 2.0%.  100 90 80 70 60  P3 S P2  50  EB/S  40  Mg (%)  P4 S  EB/S  B  EB  30  P5 S  P6 S  S EB/S EB/S  P7 S  P8 S  EB/S  EB/S EB  EB  EB  B  EB  EB  Cm(%)  20  0 0.5 1.0 2.0 3.0  DB  Impeller  10 0  5  10  15  20  25  x/D  Figure 7.6: Gas mixing index and flow regimes as a function of dimensionless distance downstream for Usl = 1.0 m/s, Usg = 0.11 m/s, N = 600 rpm and various fibre mass concentrations.  193  7.3.3 Effect of superficial gas velocity on gas-suspension mixing Figure 7.7 illustrates the influence of superficial gas velocity on mixing for Usl = 0.5 m/s and almost identical rotational speeds. At Cm = 1.0%, buoyancy caused the gas to segregate immediately downstream of the impeller, rapidly approaching stratified flow, as shown in Figure 7.7a. Mixing was almost independent of the gas velocity, and Mg reached a similar value of ~0.7 as for air-water flow. The influence of superficial gas velocity on mixing at a higher concentration, Cm = 2.0%, is shown in Figure 7.7b. The robust fibre networks deflected liquid/pulp slugs to the top of the pipe, resulting in elongated bubble flow downstream of P2 for Usg ≤ 0.22 m/s, whereas stratified flow occurred at P3 for Usg ≥ 0.33 m/s, with elongated bubble flow downstream. Mixing deteriorated with increasing superficial gas velocity, due to larger bubbles and lower slug frequency. Figure 7.8 compares mixing with and without the impeller for Cm = 1.0%, Usl = 1.0 m/s and different superficial gas velocities. At Usg = 0.055 m/s, the impeller improved gas dispersion slightly in the P2 plane, with mixing quality approaching that without the impeller downstream for elongated bubble flow, as shown in Figure 7.8a. The impeller, however, worsened mixing considerably at Usg = 0.11 m/s, corresponding to gas volume fraction of ~0.1, as shown in Figure 7.8b, especially in the high-shear (P2) zone, since a gas cavity formed around the impeller. Buoyancy also caused gas to segregate to the top of the pipe (as noted when discussing Figure 7.6), leading to stratified flow and poor mixing downstream.  194  (a) 100 90  Mg (%)  80 70  EB/S  60  P2  S  S P3  P4  S  S  S  S  P5  P6  P7  P8  Usg(m/s) N (rpm)  50  0.11 0.22 0.33 0.44  40  437 433 440 432  Impeller 30 0  5  10  15  20  25  x/D  (b) 100 90 80 70 60  S  EB/S  EB  50 EB  Mg (%)  40  EB  30  EB EB  EB  EB  EB  EB  EB  EB  EB  EB  EB  EB EB EB  EB EB  EB  Usg(m/s) N (rpm)  20  0.11 0.22 0.33 0.44  Impeller  421 418 411 413  10 0  5  10  15  20  25  x/D  Figure 7.7: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Usl = 0.5 m/s, virtually identical impeller speeds and various superficial gas velocities at: (a) Cm = 1.0%; (b) Cm = 2.0%.  195  (a) 100  Mg (%)  B/EB  EB EB  EB  EB  EB  DB 10  Usg (m/s) N (rpm)  DB  0.055 0.055  Impeller 5  0  5  10  NA 440  15  20  25  x/D  (b) 100 90 80 70 60  S  S  B/EB  EB  S  S  EB  EB  EB  EB/S  50 40  Mg (%)  S  S  EB  30  Usg (m/s) N (rpm)  20 DB  0.11 0.11  NA 440  Impeller 10 0  5  10  15  20  25  B  Figure 7.8: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Usl = 1.0 m/s and Cm = 1.0%, with and without an impeller at: (a) Usg = 0.055 m/s; (b) Usg = 0.11 m/s.  196  Figure 7.9 compares mixing with the impeller (bold lines) and without it (dash lines) at a higher Usl, 3.0 m/s, for various superficial gas velocities. At Usg = 0.11 m/s, the impeller improved mixing substantially, both in the high-shear zone and downstream (see Figure 7.9a). The impeller efficiency decreased as more gas was present in the system at Usg = 0.33 m/s, but the improvement in mixing with the presence of the impeller was still significantly greater than at Usl of 1.0 m/s and Usg = 0.055 m/s, even though the gas volume fraction (0.1) was approximately twice that at Usl = 1.0 m/s (compare Figure 7.8a). The gas volume fraction causing the impeller to worsen mixing occurred at Usg = 0.44 m/s (gas volume fraction of 0.13), as shown in Figure 7.9b, whereas this occurred at Usg = 0.11 for Usl = 1.0 m/s (gas fraction of 0.1). For this condition (Usl = 3.0 m/s and Usg = 0.44 m/s), a gas cavity formed around the impeller, causing worse mixing in the high-shear zone, with better mixing than without the impeller downstream, whereas mixing was considerably worse than without the impeller, both in high-shear zone and along the pipe, when flooding occurred at Usl = 1.0 m/s (compare Figure 7.8b). Increasing shear by increasing Usl reduced gas formation around the impeller, increasing the gas volume fraction that can be handled by the system, in addition to enhancing downstream mixing. Figure 7.10 compares tomographic images at Usg = 0.33 and 0.44 m/s for the same operating conditions. Increasing Usg from 0.33 to 0.44 m/s (gas volume fraction from 0.1 to 0.13) caused a dramatic increase in gas holdup. At the P2 plane immediately downstream of the impeller, the gas region (red, yellow and green) was  197  (a) 100  Usg (m/s) N (rpm) 0.11 0.11 0.33 0.33  Mg (%)  B DB  NA 433 NA 428  B/EB  B/EB B  B  10  DB  Impeller DB 0  B  DB  DB  5  B  DB  B  B/EB B/EB  B  DB  B/EB  DB  DB  5  10  15  20  25  x/D  (b) 100 90 80 70 60  Usg (m/s) N (rpm)  50  0.44 NA 0.44 432 B/EB B/EB B/EB  B  Mg (%)  40  B B/EB  B  30  DB  B/EB  DB B  20 DB  B/EB  B  Impeller  10 0  5  10  15  20  25  B  Figure 7.9: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Usl = 3.0 m/s and Cm = 1.0%, with and without an impeller at: (a) Usg = 0.11 and 0.33 m/s; (b) Usg = 0.44 m/s.  198  significantly more concentrated at Usg = 0.44 m/s than at Usg = 0.33 m/s, clearly indicating flooding around the impeller. Higher shear from high superficial liquid/pulp velocity improved mixing significantly, especially at the P3 and P4 planes, with reflocculation and buoyancy causing gas to segregate and coalesce downstream at the top the pipe. The difference in gas holdup for Usg = 0.33 and 0.44 m/s for each downstream position was, however, significantly less than around the impeller. Figure 7.11 illustrates the influence of superficial gas velocity on mixing at Usl = 3.0 m/s and Cm = 3.0% and compares mixing with and without an impeller. At this highest superficial liquid/pulp velocity and fibre mass concentration investigated, mixing in the high shear (P2) zone was substantially worse for Usg ≥ 0.33 m/s than for lower Usg, as shown in Figure 7.11a, indicating flooding around the impeller at the gas volume fraction of 0.1, lower than for lower Cm (see Figure 7.9). Below this gas volume fraction (Usg ≤ 0.22 m/s), mixing with the impeller was significantly better than without it (Figure 7.11b), especially in the high-shear zone, since the impeller dispersed bubbles throughout the cross-section. Beyond this gas volume fraction, mixing with the impeller was similar to that without it, whereas the impeller worsened mixing considerably at a lower Usl (1.0 m/s). Mixing deteriorated downstream of the impeller, likely because reflocculation and buoyancy caused bubbles to congregate at the top of the pipe.  199  (a) Top  (b)  Figure 7.10: Tomographic images for air-suspension flow for Usl = 3.0 m/s and Cm  = 1.0% at: (a) Usg = 0.33 m/s; (b) Usg = 0.44 m/s. Locations of planes P1 – P8 are shown in Figure 3.5. Colour scales indicate gas holdup.  200  (a) 100  EB  B  Mg (%)  B  B  10  EB  B/EB  B/EB  B  B/EB  B  DB DB  DB  B/EB  Usg (m/s) N (rpm)  B  0.11 0.22 0.33 0.44  Impeller 0  5  10  B/EB B/EB  DB 5  EB  B/EB  B  DB  EB  15  423 432 441 442 20  25  x/D  (b) 100  B  EB  B  Mg (%)  B  EB  B/EB B/EB  B B  EB  B/EB B  B  B  B/EB B/EB  B/EB  Usg(m/s) 0.11 0.22 0.33 0.44  10  5  0  5  10  15  20  25  x/D  Figure 7.11: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 3.0%, Usl = 3.0 m/s and various superficial gas velocities: (a) with impeller; (b) without impeller.  201  7.3.4 Influence of superficial liquid/pulp velocity Figure 7.12 shows the effect of superficial liquid/pulp velocity on mixing for Usg = 0.11 m/s and virtually identical impeller speeds. At Cm = 1.0%, the mixing  quality decreased with decreasing Usl for each downstream position as shown in Figure 7.12a. The results were similar to those for water (Figure 7.2) except that stratified flow was found at Usl = 1.0 m/s, with mixing quality similar to Usl = 0.5 m/s. This was likely due to buoyancy and fibre networks at Cm =1.0% causing rapid gas segregation downstream of the impeller, as discussed in section 7.3.2 (Figure 7.6). At Cm = 3.0%, mixing was less dependent on Usl for each downstream position, as illustrated in Figure 7.12b. Mixing quality at P2 was almost identical for Usl ≤ 2.0 m/s, at which stronger fibre networks and buoyancy caused the gas to segregate rapidly just downstream of the impeller, whereas bubbles were dispersed throughout the cross-section for Usl = 3.0 m/s. Re-establishment of fibre networks then likely occurred downstream, causing liquid/pulp slugs for Usl ≤ 2.0 m/s and accelerating bubble coalescence at the top of the pipe for Usl = 3.0 m/s. This led to elongated bubble flow and B/EB transitional flow, with similar mixing quality for different Usl.  202  (a) 100 S  EB/S S  S  S  S  S  Usl(m/s) N (rpm)  Mg (%)  0.5 1.0 2.0 3.0  437 423 437 433  B/EB B  DB 10  5  DB P2  DB  Impeller P3  P4  0  DB  5  DB  B/EB B  B  DB DB DB  B/EB  P8  P7  P6  P5  10  15  20  25  x/D  (b) 100  Mg (%)  EB/S S EB/S EB EB/S B/EB  EB  EB  EB  EB  EB  EB  EB EB EB B  DB  DB  10  5  10  EB EB  EB EB  B/EB B/EB 0.5 1.0 2.0 3.0  Impeller 0  EB  Usl(m/s) N (rpm)  B  DB  5  EB  15  427 423 418 423 20  25  x/D  Figure 7.12: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Usg = 0.11 m/s, virtually identical impeller speeds and various superficial liquid/pulp velocities: (a) Cm = 1.0%; (b) Cm = 3.0%.  203  7.3.5 Effect of impeller speed on gas-suspension mixing Figure 7.13 plots the gas mixing index for Usl = 3.0 m/s, Usg = 0.11 m/s, Cm = 3.0% and various rotational speeds. For suspension mixing, gas uniformity in the high-shear (P2) zone was substantially better than for the tee mixer alone, whereas in air-water flow, mixing quality in the presence of the mechanical mixer improved only slightly (compare Figure 7.3a). Without an impeller, robust fibre networks caused bubbles to congregate near the bottom of the pipe, as illustrated by the high gas-concentration regions for P2 – P4 in Figure 7.14a, whereas the gas was less concentrated and distributed more uniformly for the impeller rotating at 420 rpm, as shown in Figure 7.14b. At this fibre mass concentration, the flow was essentially plug before passing through the impeller, as described in section 4.3.2. Shear from the impeller disrupted the fibre networks and dispersed bubbles throughout the cross-section, resulting in much better mixing than for the tee alone. In the latter case, buoyancy played a stronger role downstream of plane P5, causing the gas to reach to the top of the pipe, as shown by higher gas-concentration regions in Figure 7.14a. The gas then formed larger bubbles, approaching elongated bubble flow. With the mechanical mixer in operation, decaying turbulence and reflocculation likely occurred downstream, accelerating bubble coalescence at the top of the pipe, as illustrated by higher gas-concentration regions at the top downstream of P5 in Figure 7.14b, with mixing quality reaching that for the tee alone as elongated bubble flow was approached. Mixing improved only slightly in the high-shear zone as the impeller speed increased, and became almost independent of rotational speed downstream of P2. 204  100  N (rpm)  B  Mg (%)  B  NA 423 600 803 B  B  B B  B DB DB  DB  10  Impeller 0  B/EB  B/EB B/EB P8 P7  P5 P4  P2 5  P6  B/EB  P3  5  10  15  20  25  x/D  Figure 7.13: Gas mixing index and flow regimes as a function of dimensionless  distance downstream in air-suspension flow for Usl = 3.0 m/s, Usg = 0.11 m/s, Cm = 3.0% and various impeller speeds. Figure 7.15 portrays the influence of rotational speed on mixing for Usl = 1.0 m/s and Usg = 0.11 m/s. At Cm = 0.5%, increasing impeller speed beyond 600 rpm provided better mixing in the high-shear zone, as illustrated in Figure 7.15a, with improved mixing relative to higher Usl (compare Figure 7.13), suggesting that the residence time was significant and mixing could be enhanced by increasing impeller speed when the gas flow was insufficient to cause a gas cavity around the impeller. Mixing was, however, almost independent of impeller speed downstream. Higher Cm of 2.0% and the same gas volume fraction of 0.1 led to gas cavity formation  around the impeller, with buoyancy and dense fibre networks causing the gas to segregate to the top of the pipe immediately downstream of the impeller, causing mixing to be considerably worse than without the impeller (see Figure 7.15b). The influence of impeller speed on downstream mixing is also clear. For the tee mixer 205  (a)  Top  (b)  Figure 7.14: Tomographic images for air-suspension flow for Usl = 3.0 m/s, Usg =  0.11 m/s and Cm = 3.0%: (a) tee mixer alone; (b) impeller rotating at 423 rpm. Locations of planes P1 – P8 are shown in Figure 3.5. Colour scales indicate gas holdup. alone, robust fibre networks confined the gas to the top of the pipe, causing liquid/pulp slugs to flow along the upper part, leading to elongated bubble flow. The 206  impeller disintegrated the fibre networks, causing the gas to segregate to the top of the pipe without slugs, resulting in stratified flow. The rotating at 800 rpm and the perpendicular static impeller provided poor mixing downstream, since stratified flow occurred. As the impeller speed decreased, the formation of liquid/pulp slugs caused by re-establishment of fibre networks likely occurred more rapidly, as illustrated by improved mixing downstream of P5 for N = 400 and 600 rpm. For the parallel static impeller, reflocculation likely occurred rapidly downstream of P3, with mixing quality approaching that for the tee alone for elongated bubble flow. Figure 7.16 shows the influence of impeller speed and configuration on mixing for Usl = 0.5 m/s, Usg = 0.11 m/s, Cm = 3.0%. At this lowest superficial liquid/pulp velocity and the highest fibre mass concentration investigated, elongated bubble flow occurred more rapidly than for the other flow conditions tested. For this buoyancy-dominated condition, mixing in the high-shear (P2) zone improved as the rotational speed decreased, likely due to more rapid reflocculation. The stationary impeller when perpendicular to the flow disrupted fibre networks, leading to stratified flow at P2 and P3. Re-establishment of fibre networks then caused liquid/pulp slugs at the top of the pipe downstream of P3, with considerably improved mixing. Reflocculation likely occurred rapidly downstream of the impeller for N ≤ 600 rpm and for the parallel static impeller, with slugs forming immediately and mixing quality approaching that for the tee mixer alone. Supplementary data relevant to this chapter are provided in Appendix D.  207  (a) 100 90 80 70 60 50  B  EB/S  EB/S  EB/S  EB/S  EB/S  Mg (%)  40 30  N (rpm)  DB  NA 404 602 803  20  Impeller  10 0  5  10  15  20  25  x/D  (b) 100 90 80 70 60  P3 P2  S  50  P5  P6  P7  S  S  S  S  S  S  40  Mg (%)  P4 S  S S  EB/S  EB  EB/S EB/S  EB/S  30  P8 S  EB  EB  N (rpm) NA 0 (//) 0( ) 401 600 803  Impeller 10 0  5  EB  EB  B  20  EB  10  15  20  25  x/D  Figure 7.15: Gas mixing index and flow regimes as a function of dimensionless  distance downstream in air-suspension flow for Usl = 1.0 m/s, Usg = 0.11 m/s and various impeller speeds at: (a) Cm = 0.5%; (b) Cm = 2.0%.  208  100 90 80 70 60  S  S S  50  EB/S  Mg (%)  40  EB EB/S  30  P2  EB P3  EB P4  EB P5  EB  EB P6  P7  NA 0 (//) 0( ) 427 600 802  20  Impeller 10 0  5  P8  N (rpm)  10  15  20  25  x/D  Figure 7.16: Gas mixing index and flow regimes as a function of dimensionless  distance downstream in air-suspension flow for Usl = 0.5 m/s, Usg = 0.11 m/s, Cm = 3.0% and various impeller speeds.  7.4 Conclusions Gas mixing in horizontal water flow with an in-line mechanical mixer deteriorated with increasing superficial gas velocity due to larger bubbles. With sufficient gas flow, the impeller worsened mixing since a gas cavity formed around the impeller. Mixing quality depended greatly on superficial liquid velocity and flow regime. For Usl = 3.0 m/s, mixing with the impeller was much better than without it, whereas for buoyancy-dominated flow at Usl = 0.5 m/s, formation of a gas cavity around the impeller and buoyancy caused more rapid gas segregation and worse mixing than for a tee alone.  209  Buoyancy played an important role in determining gas dispersion. Fibre mass concentration, Cm, influenced mixing in a complex manner. At Usl = 3.0 m/s, gas uniformity was almost independent of Cm in the high-shear zone just downstream of the impeller, but deteriorated substantially downstream with increasing Cm. At Usl = 0.5 m/s, gas segregated rapidly to the top of the pipe due to buoyancy, leading to stratified flow for Cm ≤ 1.0%, whereas mixing improved significantly for Cm ≥ 2.0, since re-forming of robust fibre networks led to liquid/pulp slugs at the top of the pipe, resulting in elongated bubble flow. Gas volume fraction was another key factor. Formation of a gas cavity around the impeller, i.e. “flooding”, occurred when sufficient gas was added to the system, with this amount increasing with increasing the superficial liquid/pulp velocity due to higher shear reducing gas accumulation around the impeller, and decreasing with increasing fibre mass concentration due to higher suspension yield stress preventing contact between suspension and impeller. When a gas cavity formed around the impeller, mixing was substantially worse than without the impeller for buoyancy-dominated flow (Usl = 1.0 m/s), both in the high-shear zone and downstream. When buoyancy was less significant (Usl = 3.0 m/s), flooding also caused poor mixing around the impeller, but higher shear improved mixing downstream significantly compared to lower Usl. Below the gas volume fraction corresponding to cavity formation, mixing with the impeller in the high-shear zone at Usl = 3.0 m/s was significantly better than without it, but deteriorated rapidly downstream, likely due to reflocculation promoting bubble coalescence. Mixing was almost independent of impeller speed at this Usl.  210  The influence of impeller speed on mixing was more significant at longer residence times. The presence of the impeller and increasing impeller speed improved mixing at lower Usl where the gas volume fraction was insufficient for flooding. When flooding occurred, mixing was worse in the presence of the impeller and almost independent of impeller speed, with buoyancy causing the gas to segregate to the top of the pipe, leading to stratified flow. Downstream mixing, however, depended strongly on the impeller speed. Lower speed likely caused more rapid reestablishment of fibre networks, resulting in elongated bubble flow and improved mixing.  211  8. Mixing Quality for an Industrial Pulp Mixer using Electrical Resistance Tomography9  8.1 Introduction As in other industries, assessing mixing quality in pulp and paper processes is a difficult task. Several methods have been applied. Among the most common are indirect measurement techniques based on the final product quality such as kappa number (lignin content), pulp brightness, pulp cleanliness, and pulp strength (Atkinson and Partridge, 1966; Bennington et al., 2001). Poor product quality can be an indication of poor mixing in the production process. However, product quality is an ambiguous indicator of mixing quality since it is unable to specify where the problem occurs. For example, poor product quality could be due to poor washing of pulp rather than poor mixing. Therefore, direct methods to assess mixing quality are highly desirable. There are a number of direct methods to evaluate mixing quality. At a laboratory or pilot scale, these include measurement of the distribution of chemical residuals (Paterson and Kerekes, 1985), inert tracers (Breed, 1985; Bennington et al., 1997; Kamal and Bennington, 2000) and radioactive isotopes (Kuoppamaki, 1985; Kuoppamaki et al., 1992). For mill applications, the technique must not  9  A version of this chapter was published as Yenjaichon, W., Pageau G., Bhole, M., Bennington,  C.P.J, Grace, J.R, 2011. Assessment of mixing quality for an industrial pulp mixer using electrical resistance tomography. Can. J. Chem. Eng. 89(5), 996–1004.  212  interfere with the process or adversely affect pulp quality. Past techniques include measurement of the uniformity of chemical residuals (Paterson and Kerekes, 1986), distribution of inert tracers (Torregrossa, 1983; Backlund et al., 1987), and temperature profiles around process piping (Torregrossa, 1983; Sinn, 1984; Pattyson, 1984; Robitaille, 1987; Rewatkar et al., 2002). Details of these previous studies are provided in sections 2.6 and 2.7. There is still no effective method to determine mixing quality on the industrial scale. Most mixing assessment techniques are intrusive, tedious and time consuming. Since they can only give an indication at a sampling point, a large number of samples at different locations are required to give a clear picture of how mixing is occurring. Temperature profiling of the surface of a process pipe can be used to infer the mixing quality if a sufficient temperature difference exists between the pulp stream and the added chemical. It does not interfere with the process and can be implemented as an in-situ measurement technique. However, this technique only determines mixture quality at the periphery of the pipe and is unable to quantify mixing throughout the suspension. The main goals of this chapter are to develop protocols for measuring mixing quality in pulp suspensions and to implement the technique as a real-time mixing assessment tool on an industrial scale. Electrical resistance tomography has been applied to evaluate the mixing quality of pulp suspensions downstream of a static mixer in the first chlorine dioxide bleaching (D0) stage at Howe Sound Pulp and Paper Ltd. (HSPP), Port Mellon, BC for various operating conditions. The D0 stage operates with low-consistency pulp suspensions, Cm = 3 – 4%, treated with chlorine dioxide, ClO2 (1.5 – 2.0 wt.% on  213  pulp) for 20 minutes at temperatures from 50 – 60oC. These conditions are referred to as normal operating conditions in this chapter. Bleaching begins by contacting pulp with ClO2 in a brief (1 – 2 s) period inside a static mixer. The pulp is then transferred to a reaction tower with sufficient retention time to complete the reaction. The process flow diagram for the Do stage is illustrated in Figure 8.1. The ClO2 solution (10 g/L) is introduced in the pulp suspension at a T-junction through a small drilled pipe sparger that distributes the chemical throughout a cross-section of the static mixer. The mixture then flows through the static mixer to contact the pulp suspension with the added chemical and to ensure radial uniformity of composition. The static mixer is important to ensure efficient contacting between chemical and pulp in the tower, an essential requirement for achieving the greatest lignin removal and optimal use of the bleaching chemical.  Figure 8.1: Schematic of D0 stage at Howe Sound Pulp and Paper.  214  8.2 Experimental Details The ERT technique was applied to assess the quality of mixing of pulp suspension with chlorine dioxide by measuring the distribution of electrical conductivity in a cross-section of the pipe. A 610 mm ID PTFE gasket was inserted between the flanges after the static mixer. It served as an ERT measurement plane as shown in Figure 8.2. The measurement plane consists of 16 circular titanium electrodes (38.1 mm diameter), equi-spaced at 22.5o intervals around the periphery. Each electrode (sensor) was threaded into the PTFE gasket to be flush with the internal wall of the plane. Thus, the electrodes did not alter the flow of pulp suspension in the pipe, and the measurement technique was non-intrusive. A screw was tapped into the end of each electrode to attach the coaxial cables, which carried input and output electrical ERT signals. Details of the sensor design are shown in Figure 8.3. The sensors were arranged around the pipe periphery, with the first electrode at the top of the pipe. A ground electrode of the same size as the sensor electrode was positioned between electrodes 8 and 9 at the bottom of the sensor plane. All electrodes were connected to an ITS Z8000 system (Industrial Tomography Systems, Manchester, UK) via 2.5 m long coaxial cables. The ITS Z8000 system applies a constant AC current to a pair of electrodes and measures the voltage differences between the other electrode pairs using an adjacent-pair strategy (ITS, 2007). The frequency of the injection current was 80 kHz. The sampling interval was maintained at 100 ms, and hence data were acquired at 10 Hz. A linear back projection algorithm was employed for image  215  (a)  (b)  Figure 8.2: ERT sensor plane (a) located at outlet of static mixer where pulp  suspensions flow from right to left; (b) inserted between flanges downstream of static mixer. 216  (a)  (b)  Figure 8.3: ERT sensor design details (a) schematic of ERT sensor; (b) image of  electrode.  217  reconstruction using ITS Z8000 software. In general, the Z8000 system has excellent temporal resolution (maximum sampling rate of 1,000 Hz), but achieves a spatial resolution of only 5 – 10% of the pipe diameter (equivalent to 30 - 60 mm in this study). The local conductivity values from ERT measurement were temporally averaged for 100 - 150 frames (representing 10 - 15 s of operation) for further analysis. The degree of uniformity in a measurement plane as a function of time was characterized by the mixing index (M), the coefficient of variation of the electrical conductivity value in each image pixel defined by equations 3.1 and 3.2 in section 3.3.1. The modified mixing index, M’, is not used since the work in this chapter was done at an early stage before we began to use M’. In addition, as in Table 8.1, it allows comparison with earlier studies, also using M to quantify mixing.  8.3 Results and Discussion 8.3.1 New mixer installation at HSPP Komax static mixers have been used in the D0 stage at HSPP since 1989. Increasing chemical usage through 2007 and early 2008 led the mill to inspect the old mixer, and its performance was found to be unsatisfactory. It was replaced by a new static mixer during the October 2008 mill maintenance shutdown. At this time, it was found that mixer elements in the old mixer were completely missing as shown in Figure 8.4. The new Komax static mixer consists of four mixer elements with adjacent mixing elements rotated by 90. It was designed to improve the 218  performance by reducing the diameter from 610 to 508 mm to give higher velocities and hence higher energy for mixing from a higher pressure drop. During the mixer installation period, we had the opportunity to install the ERT sensor plane between flanges immediately downstream of the mixer. This allowed the performance of the new mixer to be assessed based on the new measurement technique and the results to be compared with mixing measurements which had utilized other techniques in previous studies for similar mixer installations. The chemical use before and after mixer replacement was compared based on the kappa factor (ratio of the amount of chemical applied to the process to the amount of lignin in the pulp entering the process). The kappa factor was found to significantly decrease from 0.29 to 0.24, as shown in Figure 8.5, resulting in chemical savings of approximately $1,600,000 per year. Comparison indicated that the mixing quality improved significantly when the new mixer was installed. The variation of kappa factor in Figure 8.5 is quite large on a daily basis, mainly due to variations in process conditions. The chemical use is strongly affected by the performance of brownstock washers ahead of the static mixer. The evaporator sometimes reached its upper limit, so that there was not enough fresh water to properly wash pulp, resulting in higher chemical demand. In addition, chemical use depends on the composition of the wood entering the process, and this changed during some periods of time. However, this variation did not affect the ERT measurements. Each ERT measurement period lasted less than 5 minutes, while the time scale for variation of kappa factor was about a day. Hence the impact of this variation on the ERT measurements was very small.  219  (a)  (b)  Figure 8.4: Photographs of (a) old static mixer; (b) its mixer elements.  220  Figure 8.5: Kappa factor representing chemical use before and after mixer  installation.  8.3.2 Reaction of ClO2 with pulp Chlorine dioxide added to the pulp for bleaching operation can itself be treated as a tracer in the mixing studies carried out in this work. This is possible because the electrical conductivity of ClO2 (1.76 mS/cm) differs significantly from the conductivity of pulp suspension (4.8 mS/cm). In a typical conductivity tomogram, the regions of low conductivity thus correspond to ClO2-rich regions. However, the extent of reaction needs to be considered since the tracer reacts with the process fluid. Chlorine dioxide delignification is a combination of an initial fast reaction and a 221  subsequent slow reaction (Tessier and Savoie, 1997; Chandranupap and Nguyen, 1998). The pulp suspension and chlorine dioxide are in contact for approximately 2 s (based on a typical flow velocity of 1.5 m/s in the 508 mm ID pipe corresponding to a production rate of 1,000 t/d or a suspension flow rate of 300 L/s) when the mixture arrives at the ERT measurement plane. If there were to be plug flow of the mixture through the pipe, the extent of reaction is estimated to be about 20% at the measurement plane, based on reaction kinetics of pulp suspension with chlorine dioxide (Chandranupap and Nguyen, 1998). The measured ClO2 concentration at a distance of approximately 12 m downstream of chemical injection (corresponding to a residence time of 10 s) also showed high chemical consumption, roughly 80%. Unfortunately, there is no sampling port for taking samples immediately downstream of the static mixer, but the extent of reaction is expected to be significant at the measurement plane. The reaction of a tracer with pulp reduces the tracer concentration, possibly changing the chemical distribution in the pulp suspension and thus the measured mixing quality in the process. The reaction also changes the electrical conductivity of the mixture. The conductivity of the mixture approximately 12 m downstream of chemical injection was found to be 10% higher than that of the pulp suspension before the mixer. The change in mixture conductivity due to reaction therefore does not affect the interpretation of the regions of low conductivity, which can still be referred to as ClO2-rich regions.  222  8.3.3 Variability of conductivity with time for pulp suspension and ClO2 feed to mixer Figure 8.6 illustrates the temporal variation of electrical conductivity measured 60 m upstream of the static mixer for pulp suspension and also in the ClO2 pipe upstream of injection into the pulp suspension. There was little variability of conductivity in a period of about an hour for both fluids, while the ERT measurement period was less than 5 minutes. This ensures that the errors in ERT mixing quality measurement caused by variation in the electrical conductivity of the feeds are negligible.  Figure 8.6: Temporal variation of electrical conductivity of pulp ahead of mixer and  ClO2 solution.  223  8.3.4 Flow regime For pulp suspension flow through pipes, it is customary to identify the flow regime as plug, mixed or turbulent based on the criterion given by Robertson and Mason (1957), with details provided in section 4.3.2. The suspension flow before entering the static mixer varied from 0.7 to 1.3 m/s and mass concentration from 3.2 to 4.7%, indicating that the pulp suspension flow was likely in the plug flow regime.  8.3.5 Mixing quality as a function of process operating conditions The mixing index, M, in the absence of ClO2, designated Ms, was found to be 2.87%. Figure 8.7 illustrates the measured mixing index and corrected mixing index when the ClO2 valve was opened from the shut-off position. The ClO2 valve had been closed for approximately 20 s before being opened to the normal position. The corrected mixing index significantly increased after the ClO2 valve was opened, with the average mixing index when ClO2 was present being 4.86%, which is considered to be the mixing index for normal operation. Tomographic images for the same condition are shown in Figure 8.8. As noted in section 3.3, the top position of the pipe in all the tomographic images in this chapter were rotated counterclockwise by 90 from the vertical axis due to misalignment between the electrode position in the ERT reconstruction process and the actual electrode position in the sensor plane. The first image illustrates the 224  temporal-averaged conductivity tomograph over the first 20 s with no ClO2 present, whereas the other image shows the time-average conductivity distribution with ClO2 present. The second image shows the low-conductivity region on the right of the image due to the presence of the added chemical. Although some ClO2 was likely consumed by the initial fast reaction, the amount of tracer was high enough for ERT to detect the ClO2-rich region. However, the conductivity of this region (4.0 – 4.6 mS/cm) is higher than the measured ClO2 conductivity (1.76 mS/cm) and close to the conductivity of pulp suspension (4.8 mS/cm), mainly due to the dilution of ClO2 in pulp suspensions.  Figure 8.7: Mixing index as a function of time after ClO2 valve was opened from  shut-off position.  225  Figure 8.8: Tomographic images before and after ClO2 was introduced.  Figure 8.9 shows the effect of a step change in suspension flow rate on mixing quality of the pulp suspension at a constant mass concentration of 3.5%. The suspension flow rate decreased from 390 to 225 L/s, and the mixing index increased continuously as the flow rate decreased. The reduction in mixing quality was due to the lower energy for mixing when the flow through the static mixer decreased. Mixing in a static mixer relies on the flow rate through it for its effectiveness as a mixer, specifically the energy derived from the pressure drop across the mixing elements. Thus, a sufficient flow rate must be maintained to obtain good mixing in a static mixer.  226  Figure 8.9: Temporal variation of mixing index as suspension flow rate changed  from 390 to 225 L/s at Cm = 3.5%.  The tomographic images for the same condition are illustrated in Figure 8.10. The images are in time series, each averaged over 10 s. The images show the lower conductivity region at the bottom of the image (blue in the colour reproductions), which is probably the region of high concentration of ClO2, occurring when the mixing quality decreases at lower suspension flow rate. The highconductivity region (yellow and red in the colour images) probably corresponds to the product of the reaction, consistent with the observed increase in size of this region as the suspension flow rate decreases, since one would expect an increase in the extent of reaction with increasing residence time of the suspension in the  227  pipe. The reason for the side-to-side asymmetry in this region is unclear and needs further investigation.  Figure 8.10: Tomographic images in time series when the suspension flow rate  decreased from 390 to 264 L/s at Cm = 3.5%. Figure 8.11 illustrates the change in mixing quality of the pulp suspensions when the consistency decreased from 4.7 to 3.2% at a constant suspension flow rate of 325 L/s. The mixing index was found to decrease (and hence the quality of mixing improved) with a decrease in suspension consistency, likely due to a less densely packed fibre networks, enhancing the distribution of the chemical. The region of low conductivity (blue in the colour images), likely the ClO2-rich region, was also smaller when the suspension consistency decreased as illustrated in Figure 8.12. The images were in time series with each image averaged over 15 s. The third image (t = 45 s) illustrates a decrease in the low-conductivity region, i.e. 228  an increase in the degree of uniformity compared to the first and second images (t = 15, 30 s), corresponding to the trend in Figure 8.11 when the consistency began to decrease significantly.  Figure 8.11: Temporal variation of mixing index as consistency changed from 4.7%  to 3.2% at a constant suspension flow rate of 325 L/s (Up = 1.1 m/s).  Table 8.1 summarizes the assessment of mixing quality of industrial pulp mixers based on various measurement techniques. Each of these studies employed the coefficient of variation to quantify the degree of mixing. The flow regime in each work has also been estimated based on the data available in the literature. For most of the previous work cited, the flow was in the plug flow regime, except for the study which utilized the radioactive tracer technique (Kuoppamaki, 1985) where relevant 229  data are unavailable. The mixing indices were compared, and good agreement is observed between the ERT measurement and other techniques. The mixing index measured by the ERT technique for normal operation was 4.86%, showing that good mixing quality was obtained, and this corresponded to a significant reduction in chemical use after the new mixer was installed. The good agreement with other techniques also indicates that the reactive tracer can be effective and that the extent of reaction did not significantly affect the measured mixing quality. The results also show that ERT was reasonably successful in monitoring the industrial process changes and in determining the mixer performance. The ERT technique has advantages relative to other techniques since it is non-intrusive, not tedious, and able to quantify the mixing throughout the entire suspension volume. This technique therefore has potential as an on-line, real-time mixing assessment tool for industrial pulp mills.  Figure 8.12: Tomographic Images over time as the consistency decreased from  4.7% to 3.4% at a constant suspension flow rate of 325 L/s. 230  Table 8.1: Comparison of mixing assessment of industrial pulp mixers Mixer Type Static  High Shear  Peg  Consistency, Cm (%) Lowa 2.5 3.5 9.5 -12 10.5 - 12 10 Mediumb 8 - 11 Mediumb  Measurement Technique Radioactive Tracer Chemical Residuals ERT Inert Tracer (LiCl) Temperature Inert Tracer (LiCl) Temperature Temperature Temperature  Mixing Index, M (%) 0.5 – 12 5.3 4.86 2–5 8.5 6–8 2.4 10 -40 5.8  a  Low-consistency application (Cm < 5%)  b  Medium-consistency application (8% < Cm < 16%)  Reference Kuoppamaki (1985) Paterson and Kerekes (1986) This work Kolmodin (1984) Pattyson (1984) Backlund et al. (1987) Robitaille (1987) Rewatkar et al. (2001) Robitaille (1987)  While ERT has several benefits over other techniques, it also has some drawbacks. ERT has a spatial resolution of only 5 – 10% of the pipe diameter; thus it provides limited accuracy of measurement on an industrial scale where the pipe diameter can be large. In addition, a custom sensor array has to be built, and there are constraints on placement since the installation must be between flanges. Once the sensor array is installed, it cannot be easily moved to test other mixers. The array in our case was installed between flanges downstream of the mixer, and closer to the mixer than would be optimal since the turbulence from the last mixing element continues to cause mixing for several pipe diameters downstream of the mixer exit. However, turbulence generated in the static mixer decays very quickly due to the network strength of pulp fibre suspensions, when the power dissipation required to sustain the turbulence is not sustained (Bennington et al., 1989). Therefore, mixing occurring downstream of the last mixing element of static mixer is  231  expected to be very small. Finally, this technique requires a significant electrical conductivity difference between the fluids being mixed.  8.4 Conclusions Electrical resistance tomography (ERT) was successful in assessing the mixing performance of a static mixer and in monitoring process changes in the first chlorine dioxide bleaching (D0) stage at Howe Sound Pulp and Paper. ERT was able to detect the presence of ClO2 by showing the regions of lower conductivity in the tomogram when ClO2 was introduced. The mixing quality was quantified from image pixels, and the results showed higher mixing indices (lower-quality mixing) when the ClO2 flow rate increased. In addition, ERT was able to monitor the changes in mixing quality as the suspension flow rate or mass concentration changed. At a lower suspension flow rate, the tomographic images showed regions of lower conductivity at the bottom of the pipe, indicating a higher concentration of ClO2 at that location. This was likely due to poor mixing when the energy for mixing decreased at a lower suspension flow rate. The images also showed smaller regions of low conductivity or lower concentration of ClO2 when the suspension mass concentration decreased, likely due to a decrease in the network strength of fibre suspension improving chemical distribution. In addition, quantitative analysis showed an increase in the mixing index when the suspension flow rate decreased, and a decrease in the mixing index when the suspension mass concentration  232  decreased. This demonstrates that better mixing performance can be obtained at a higher suspension flow rate and at a lower suspension consistency. The mixing index of the static mixer under normal operation obtained by the ERT technique was in good agreement with results obtained by other techniques for similar industrial pulp mixers.  233  9. Conclusions, Contributions and Recommendations 9.1 Conclusions 9.1.1  Electrical  resistance  tomography  as  a  measurement  technique Electrical resistance tomography was successfully implemented to assess the mixing performance of a static mixer and to monitor the changes in the mixing quality as a function of process operating conditions in an industrial pulp bleaching stage. ERT was able to measure mixing quantitatively, indicating lower mixing quality as the chemical flow rate increased and when the suspension flow rate decreased. It also showed that mixing improved with decreasing suspension mass concentration at a constant volumetric suspension flow rate. This demonstrates that sufficient suspension flow rates must be maintained to obtain good mixing, and that better mixing can be achieved at higher suspension flow rates, as well as at lower suspension concentrations. The results are in good agreement with those in the literature based on other measurement techniques for similar mixer installations. The ERT technique has advantages over other techniques since it is non-intrusive, not tedious, and able to quantify the mixing throughout the entire suspension volume. However, this technique has a spatial resolution of only 5 – 10% of the pipe diameter, one or more custom sensor array has to be installed, and a significant electrical conductivity difference is required between the fluids being mixed.  234  9.1.2 Liquid jet mixing into liquid mainstream In-line jet mixing in pulp suspensions depended strongly on the flow regime and jet penetration. For turbulent flow, the criteria for in-line jet mixing in water apply also to the mixing in suspensions. Mixing improved considerably with increasing jetto-pipe velocity ratio, but was almost independent of the mainstream velocity for a given jet-to-pipe diameter ratio and the conditions studied. Small differences in mixing quality for water and turbulent suspension flow were likely due to differences in fibre network strength and influences of fibre-turbulence interactions in modifying turbulent structures in the bulk. For the mixed flow regime with loose fibre networks, mixing improved significantly with increasing mainstream velocity, whereas mixing was poor at higher mass concentrations (Cm ≥ 2.0% for softwood pulp and Cm = 3.0% for hardwood pulp) for plug flow, even at high mainstream velocities, since shear was insufficient to disrupt the plug. For plug flow in the mixed flow regime with dilute suspensions, mixing improved with increasing jet velocity since the jet was able to disrupt the plug. At higher fibre mass concentrations (Cm ≥ 1.0% with softwood pulp and Cm ≥ 2.0% with hardwood pulp), jet penetration played an important role. Mixing was poor when the jet attached to the far wall of the pipe, but improved when the jet penetrated to the axis of the pipe, and was even better when the jet impinged on the opposite wall and then recirculated to the core of the pipe. Velocities required for the jet to penetrate to the centre of the pipe were higher at higher mass concentrations and lower mainstream velocities.  235  Mixing with an in-line mechanical mixer was mostly similar to that for a tee mixer except that: (a) for Cm ≥ 2.0%, mixing with the impeller improved significantly with increasing mainstream velocity, whereas it was only slightly better for no impeller, and (b) mixing with the impeller was worse when the jet penetrated to the far wall of the pipe ahead of the impeller for all fibre mass concentrations investigated. Mixing improved considerably with increasing impeller speed at a low mainstream velocity, whereas the rotational speed had less influence at higher velocities (Up ≥ 2.0 m/s). Residence time also had more effect on mixing at higher impeller speeds. With an impeller present, mixing quality improved substantially and was similar to that for water when the flow approached the turbulent regime, with a considerably lower mainstream velocity required compared to a tee mixer alone. At higher mass concentrations, the energy from the impeller and the flow was insufficient to provide the same level of turbulence as in water, even at the highest mainstream velocity and impeller rotation speed examined. Improved mixing with increasing impeller speed primarily occurred in the high-shear zone around the impeller, with rapid decaying turbulence accompanied by reflocculation, reducing the quality of mixing downstream. Further increasing the impeller speed would likely “fluidize” suspensions in the high-shear zone, providing a similar level of turbulence and mixing quality as for water. For the range of the operating conditions investigated, optimum liquid mixing was obtained at the highest mainstream velocity and jet velocity for the jet to penetrate to the centre of the pipe, almost independent of the impeller speed. Increasing impeller speed at lower mainstream velocities (Up  236  ≤ 1.0 m/s), however, enhanced mixing in the high-shear zone, approaching that at the highest mainstream velocity, with no improvement of mixing downstream. Further increasing the impeller speed beyond the conditions in the present work would likely provide better mixing than at higher velocities in the high-shear zone, and the downstream mixing would also probably be better than at higher velocities for high impeller speeds. For industrial applications, where the fibre mass concentration is much higher than those investigated in this study, in-line mechanical mixing that creates a highshear zone is essential, since shear from suspension flow alone is insufficient to improve mixing downstream, even at high mainstream velocity. Mechanical mixing relies on the impeller speed and residence time in the high-shear zone. Thus, such mixing should be operated at a low suspension flow rate for high residence time (Usl ≤ 1.0 m/s), and at a high impeller rotation speed to obtain sufficient shear to disrupt robust fibre networks.  9.1.3 Gas mixing into liquid mainstream Gas and suspension flow regimes and buoyancy play an important role on gas dispersion in horizontal pulp suspension flow. Gas mixing depended greatly on fibre mass concentration for the bubble flow regime. For Cm ≤ 1.0%, the uniformity of bubbles was similar to that for water since the flow was turbulent, but worse at higher concentrations (Cm ≥ 2.0%) for plug flow, likely because dense fibre networks in the core of the pipe caused bubbles to congregate near the wall. When buoyancy was significant, gas uniformity improved with increasing pulp concentration, since 237  robust fibre networks led to liquid/pulp slugs at the top of the pipe, whereas stratified flow was approached at lower concentrations. Mixing was less dependent on superficial liquid/pulp velocity at higher pulp concentrations, due to less variation in flow regimes. Gas mixing decreased with increasing superficial gas velocity for dispersed bubble and bubble flow, due to more bubbles coalescing at the top of the pipe, and for elongated bubble flow because of larger bubbles and lower liquid/pulp slug frequency. Mixing was almost independent of Usg for stratified flow. Gas mixing in pulp suspensions with an in-line mechanical mixer depends on fibre mass concentration, gas volume fraction, impeller speed and buoyancy in a complex manner. With sufficient flow of gas, a gas cavity formed around the impeller, leading to considerably worse mixing. The gas volume fraction that can be handled by the impeller increased with increasing superficial liquid/pulp velocity, and decreased with increasing fibre mass concentration. Below this gas volume fraction, for buoyancy-dominated flow, mixing improved slightly in the presence of the impeller and with increasing impeller speed in the high-shear zone, and was similar to that without impeller downstream, whereas the impeller improved mixing substantially when buoyancy was not significant at higher superficial liquid/pulp velocities, with mixing almost independent of the impeller speed due to the short residence time in the impeller zone. Beyond this gas volume fraction, the impeller worsened mixing considerably for buoyancy-dominated flow, both in the high-shear zone and downstream, due to gas formation around the impeller and buoyancy causing the gas to segregate to the top of the pipe, leading to stratified flow downstream. When buoyancy was less significant, mixing was also poor in the high-  238  shear zone due to flooding, but improved significantly downstream compared to that at lower Usl. Below the gas volume fraction corresponding to gas cavity formation, the impeller dispersed bubbles uniformly in the pipe cross-section when buoyancy was not significant, with more rapid reflocculation promoting bubble coalescence and demixing downstream as the suspension concentration increased. For buoyancydominated flow, flooding caused mixing to be poor in the high-shear zone, but mixing improved downstream at higher suspension concentrations, since fibre networks likely re-formed and caused liquid/pulp slugs at the top of the pipe, whereas stratified flow occurred at lower concentrations. Mixing improved with decreasing impeller speed, likely because fibre networks re-established more rapidly, leading to elongated bubble flow. For industrial applications, dispersed bubble flow is desirable. Without an impeller, this flow regime cannot be obtained, even at very high superficial liquid/pulp velocity, since gas volume fractions and suspension concentrations in industrial gas mixing operations are normally high, with robust fibre networks of a plug in the core of the pipe causing gas to congregate near the pipe wall. The impeller can also worsen mixing considerably for buoyancy-dominated flow when the gas volume fraction is sufficient for a gas cavity to form around the impeller. It is therefore recommended to operate with an impeller at high superficial liquid/pulp velocity (Usl ≥ 3.0 m/s) so that buoyancy is not significant and at low gas volume fraction (below 0.1 for the laboratory mixer and the highest Usl and Cm investigated)  239  to prevent gas cavity formation around the impeller, hence leading to dispersed bubble flow and optimum mixing.  9.2 Contributions and Potential Applications 1. In-line liquid jet mixing behaviour for non-Newtonian pulp suspensions has been investigated, leading to better understanding. Criteria for optimum mixing have been developed for softwood and hardwood fibre mass concentrations of 0.5 – 3.0%, providing basic concepts and guidance for the design and optimization of in-line mixers combined with injectors. 2. The influences of gas and suspension flow regimes and buoyancy on gas dispersion in horizontal pulp suspension flow have been addressed, leading to better knowledge of gas-fibre-suspension multiphase flow and basic concepts for the design of in-line mixers with gas injection. 3. Suspension mixing with an in-line mechanical mixer has been evaluated for various operating conditions. Optimum operating conditions have been established for both liquid-suspension and gas-suspension mixing, providing useful guidance for optimizing in-line mixers in pulping operations. 4. ERT has been implemented on an industrial scale in the pulp and paper system for the first time, with successful results in assessing the mixing performance of a static mixer and in monitoring process changes in the first chlorine dioxide bleaching (D0) stage.  240  5. Two novel mixing indices have been developed to quantify liquid-suspension and gas-suspension mixing, providing quantitative data for process optimization, with potential for application to other mixing processes.  9.3 Recommendations for Future Research 1. Basic concepts for in-line jet and mechanical mixers from this study could be applied to investigate mixing for other in-line mixers commonly used in pulp bleaching processes, such as static mixers and high-shear mixers. 2. Criteria for optimizing in-line jet mixing have only been developed for fibre mass concentration up to 3.0%, due to limitation of the laboratory flow loop, including pipe size and pump capacity. It would be worthwhile to establish the optimum jet mixing conditions at higher mass concentrations, in particular the medium-consistency range, i.e. mass concentrations from 8 to 16%. 3. The optimum operating conditions for liquid mixing with an in-line mechanical mixer have been identified based on the impeller rotational speed up to 800 rpm. Residence time was, however, found to be more significant at lower mainstream velocities (Up ≤ 1.0 m/s), with significantly improved mixing in the high-shear zone as the impeller speed increased. Operating an impeller at much higher rotational speed and low mainstream velocities would probably provide better mixing than at higher mainstream velocities. A more robust inline mixer providing higher rotational speed, such as an in-line rotor-stator mixer, could be utilized to investigate this effect.  241  4. It would be worthwhile to examine the influence of power dissipation on both liquid and gas mixing for different gas and suspension flow regimes. The mechanical mixer system could be modified to provide reliable torque readings and to determine power dissipation for different operating conditions. 5. Resistor adaptors helped resolve anomalous voltage data and reconstructed images where electrodes lose electrical contact when the gas phase concentrates at the top of a horizontal pipe. The images are, however, imperfect, due to an “edge effect”. The reconstruction process could be modified, e.g. by including a Boolean logic scheme to identify the presence of gas at the top of the pipe. 6. It would be worthwhile to test different methods to introduce gas into the system, e.g. a coaxial jet or a sparger that can produce small bubbles, if possible leading to dispersed bubble flow. 7. Other more common types of mechanical mixer impellers should be tested to determine if they are more effective than the one studied here for breaking up fibre flocs and dispersing gas into pulp suspensions. 8. 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Des. 84(A11), 993–1000.  261  Appendix A: Supplementary Data for Chapter 4 This appendix includes additional experimental results for in-line liquid jet mixing in water and pulp suspensions. 100  M' (%)  10  1  0.1 0  5  10  x/D  15  20    R  T ( C)  2.07 2.07 7.76 7.89 12.5 12.5 15.5 15.8 25.3 25.0  15.6 25.3 15.9 25.2 15.4 24.7 16.1 25.2 15.6 24.7  25  Figure A.1: Modified mixing index as a function of dimensionless distance  downstream, x/D, with various jet-to-pipe velocity ratios and mainstream temperatures for water at Dr = 0.05.  (b)  100  Up(m/s) 1.5 2.0 3.0 3.0 (w)  10  R  100 Up(m/s)  4.60 4.38 4.21 4.23  M' (%)  M' (%)  (a)  R  1.0 6.87 2.0 6.26 2.0 (w) 6.36  10  1  1 0  5  10  15  x/D  20  25  0  5  10  15  20  25  x/D  Figure A.2: Modified mixing index as a function of dimensionless distance  downstream of injection for softwood pulp suspension at Cm = 0.5% and for water (w), Dr = 0.05 with various mainstream velocities for: (a) R ≈ 4; (b) R ≈ 6.  262  Up(m/s)  100  R  1.0 2.0 3.0 4.0 4.0 (w)  10  3.99 3.46 3.20 3.21 3.17  (b) 100  Up(m/s) 1.5 2.0 3.0 3.0 (w)  10  M' (%)  M' (%)  (a)  R 4.72 4.51 4.22 4.23  1  1  0  5  10  15  20  25  0  5  10  x/D  x/D  15  20  25  Figure A.3: Modified mixing index as a function of dimensionless distance  downstream for softwood pulp suspension at Cm = 1.0% and for water (w), Dr = 0.05 with various mainstream velocities for: (a) R ≈ 3; (b) R ≈ 4. Uj (m/s)  100  10  (b)  R  3.96 6.93 9.03 12.7 12.7 (w)  Uj (m/s)  100  1.98 3.46 4.15 6.33 6.36  R  4.01 3.99 8.82 8.80 10.8 10.7 12.7 12.7 12.6 (w)12.6  10  M' (%)  M' (%)  (a)  1  1 0  5  10  x/D  15  20  0.1  25  0  5  10  x/D  15  20  25  (c)  100 Uj(m/s)  M' (%)  10  R  3.88 7.72 5.01 9.80 7.07 14.1 8.76 16.9 10.8 21.5 12.7 25.0 12.6 (w) 25.1  1  0.1 0  5  10  15  20  25  x/D  Figure A.4: Modified mixing index as a function of dimensionless distance  downstream for softwood pulp suspension at Cm = 1.0% and for water, Dr = 0.05 and various jet velocities for: (a) Up = 2.0 m/s in wall-source and jet-mixing modes; (b) Up = 1.0 m/s in jet-mixing and jet-mixing modes; (c) Up = 0.5 m/s in jet-mixing and jet-mixing modes 263  100  10  Up(m/s)  R  1.01 1.51 2.01 3.00 3.00 (w)  4.77 4.56 4.37 4.23 4.23  (b)  100  M' (%)  M' (%)  (a)  Up (m/s)  R  0.51 1.01 1.51 2.01 2.00 (w)  10  7.64 6.82 6.39 6.35 6.36  1  1 0  5  10  x/D  15  20  0  25  5  10  x/D  15  20  25  Figure A.5: Modified mixing index as a function of dimensionless distance  downstream for softwood pulp suspension at Cm = 2.0% and for water (w), Dr = 0.05 with various mainstream velocities for: (a) R ≈ 4; (b) R ≈ 6.  100  (b)  Up (m/s)  R  2.0 3.0 4.0 5.0 5.0 (w)  2.11 2.08 2.03 2.05 2.17  10  1  0  5  10  x/D  15  20  100  Up(m/s) 2.0 3.0 4.0 4.0 (w)  M' (%)  M' (%)  (a)  10  1  25  (c) 100  0  5 Up(m/s) 1.0 2.0 3.0 3.0 (w)  M' (%)  R 3.14 3.11 3.11 3.22  10  x/D  15  20  25  R 4.09 4.09 4.14 4.23  10  1 0  5  10  x/D  15  20  25  Figure A.6: Modified mixing index as a function of dimensionless distance  downstream for softwood pulp suspension at Cm = 3.0% and for water, Dr = 0.05 with various mainstream velocities for: (a) R ≈ 2; (b) R ≈ 3; (c) R ≈ 4. 264  (a)  Uj (m/s)  R  4.26 2.11 6.30 3.14 8.26 4.09 12.6 6.26 12.7 (w) 6.36  10  (b) 100  Uj (m/s)  R  4.13 4.88 6.38 8.64 10.3 12.5  8.10 9.39 12.5 16.6 19.7 24.1  M' (%)  M' (%)  100  10  1 0  5  10  x/D  15  20  1  25  0  5  10  x/D  15  20  25  Figure A.7: Modified mixing index as a function of dimensionless distance  downstream for softwood pulp suspension at Cm = 3.0% and for water, Dr = 0.05 and various jet velocities for: (a) Up = 2.0 m/s in wall-source and jet-mixing modes; (b) Up = 0.5 m/s in jet-mixing and jet-impaction modes.  100  M' (%)  Up(m/s) 1.0 2.0 3.0 3.0 (w)  10  R 4.12 4.20 4.11 4.23  1 0  5  10  15  20  25  x/D  Figure A.8: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension at Cm = 0.5% and for water, Dr = 0.05 with various mainstream velocities and similar jet-to-pipe velocity ratios of ~ 4.  265  Uj (m/s) 4.13 8.14 10.0 12.5 12.6 (w)  10  M' (%)  R 4.12 8.12 10.1 12.5 12.6  1  0.1  (b) 100  U j (m/s)  R  4.13 6.15 8.14 10.0 12.5 12.6 (w)  10  M' (%)  (a) 100  8.20 12.2 16.2 19.9 24.9 25.1  1  0.1  0  5  10  x/D  15  20  25  0  5  10  x/D  15  20  25  Figure A.9: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension at Cm = 0.5% and for water, Dr = 0.05 and various jet velocities in jet-mixing and jet-impaction modes for: (a) Up = 1.0 m/s; (b) Up = 1.0 m/s.  Up(m/s) 1.0 2.0 3.0 3.0 (w)  10  (b) 100  R 4.25 4.10 4.17 4.23  1  Up(m/s)  M' (%)  M' (%)  (a) 100  R  1.0 6.18 2.0 6.23 2.0 (w) 6.36  10  1  0  5  10  x/D  15  20  25  0  5  10  15  20  25  x/D  Figure A.10: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension at Cm = 1.0% and for water, Dr = 0.05 with various mainstream velocities for: (a) R ≈ 4; (b) R ≈ 6.  266  (b)  100  Uj (m/s)  R  4.13 6.13 9.22 12.5 12.7 (w)  10  100 Uj (m/s)  1.38 2.04 3.07 4.17 4.17  M' (%)  M' (%)  (a)  1  6.09 8.20 12.5 12.7 (w)  10  R 3.05 4.10 6.23 6.36  1  0  5  10  15  20  25  0  5  10  x/D  (c)  Uj (m/s)  15  20  25  R  4.26 4.25 6.20 6.18 8.20 8.17 10.1 10.1 12.5 12.5 12.6 (w) 12.6  100  10  M' (%)  x/D  1  0.1  0  5  10  x/D  15  20  25  Figure A.11: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension at Cm = 1.0% and for water, Dr = 0.05 and various jet velocities for: (a) Up = 3.0 m/s in wall-source and jet-mixing modes; (b) Up = 2.0 m/s in wall-source and jet-mixing modes; (c) Up = 1.0 m/s in jet-mixing and jet-impaction modes.  267  10  1  0  5  10  x/D  15  R  2.0 3.0 4.0 5.0 5.0 (w)  2.12 2.04 2.06 2.04 2.17  20 Up(m/s) 1.0 2.0 3.0 3.0 (w)  10  (b) 100  R 3.12 3.11 3.16 3.17  10  1  25 R  Up (m/s) 2.0 3.0 4.0 4.0 (w)  0  5  10  x/D  15  20  25  (d) 100 Up(m/s)  4.09 4.19 4.17 4.23  M' (%)  M' (%)  (c) 100  Up (m/s)  M' (%)  M' (%)  (a) 100  1  R  1.0 6.26 2.0 6.23 2.0 (w) 6.36  10  1  0  5  10  x/D  15  20  25  0  5  10  x/D  15  20  25  Figure A.12: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension at Cm = 2.0% and for water, Dr = 0.05 with various mainstream velocities for: (a) R ≈ 2; (b) R ≈ 3; (c) R ≈ 4; (d) R ≈ 6. 100  Uj (m/s) 6.13 9.35 12.5 12.7 (w)  10  R  (b)  2.04 3.11 4.17 4.23  M' (%)  M' (%)  (a)  1  Uj (m/s)  100  4.26 6.26 8.39 12.5 12.7 (w)  10  R 2.12 3.12 4.19 6.23 6.36  1  0  5  10  x/D  15  20  25  0  5  10  x/D  15  20  25  Figure A.13: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension at Cm = 2.0% and for water, Dr = 0.05 and various jet velocities in wall-source and jet-mixing modes for: (a) Up = 3.0 m/s; (b) Up = 2.0 m/s. 268  (a)  Uj (m/s)  100  1  0.1  (b)100  M' (%)  10  M' (%)  R  4.13 4.09 6.34 6.26 8.26 8.10 12.5 12.4 12.6 (w) 12.6  5  10  x/D  15  20  4.26 6.26 8.14 10.1 12.5  10  1  0  Uj (m/s) R  25  0  5  10  x/D  15  20  8.18 12.0 15.4 19.3 24.5  25  Figure A.14: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension at Cm = 2.0% and for water, Dr = 0.05 and various jet velocities in jet-mixing and jet-impaction modes for (a) Up = 1.0 m/s; (b) Up = 0.5 m/s.  100  Up (m/s)  R  2.0 3.0 4.0 5.0 5.0 (w)  2.0 2.06 2.06 2.05 2.17  10  (b) 100  Up(m/s) 1.0 2.0 3.0 3.0 (w)  M' (%)  M' (%)  (a)  R 4.05 4.10 4.14 4.23  10  1  1  0  5  10  x/D  15  20  25  0  5  10  x/D  15  20  25  Figure A.15: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension at Cm = 3.0% and for water, Dr = 0.05 with various mainstream velocities for: (a) R ≈ 2; (b) R ≈ 4.  269  (a)  Uj (m/s)  100  R 2.06 3.05 4.14 4.23  10  M' (%)  M' (%)  6.20 9.16 12.5 12.7 (w)  (b)  1  Uj (m/s)  100  R  4.01 6.13 8.26 12.4 12.7 (w)  2.0 3.06 4.10 6.13 6.36  10  1 0  5  10  x/D  15  20  25  0  5  10  x/D  15  20  25  (c) 100  M' (%)  10  Uj (m/s)  1  0.1  R  4.09 4.05 6.26 6.14 8.26 8.10 12.5 12.3 12.6 (w) 12.6  0  5  10  x/D  15  20  25  Figure A.16: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension at Cm = 3.0% and for water, Dr = 0.05 and various jet velocities for: (a) Up = 3.0 m/s in wall-source and jet-mixing modes; (b) Up = 2.0 m/s in wall-source and jet-mixing modes; (c) Up = 1.0 m/s in jet-mixing and jet-impaction modes  270  Appendix B: Supplementary Data for Chapter 5 This appendix includes additional experimental results for mixing liquid into water and into pulp suspensions using an in-line mechanical mixer. N (rpm) R  100  NA 0 (//) 0( ) 401 603 805  M' (%)  10 1  6.43 6.13 6.10 6.11 6.14 6.00  (b)  100  10  M' (%)  (a)  Impeller 0  5  R  NA 0 (//) 0( ) 400 600 801  1.90 2.13 2.13 2.02 2.06 2.08  1  0.1 0.01  N (rpm)  Impeller  10  x/D  15  20  0.1  25  0  5  10  15  20  25  x/D  Figure B.1: Modified mixing index as a function of dimensionless distance  downstream of injection for water for Dr = 0.05 and various rotation speeds with: (a) Up = 1.0 m/s and R ≈ 6; (b) Up = 2.0 m/s and R ≈ 2. 100  Up (m/s) R 1.0 6.26 2.0 6.20 2.0 (w) 6.20  M' (%)  10  1 Impeller 0.1  0  5  10  15  20  25  x/D  Figure B.2: Modified mixing index as a function of dimensionless distance  downstream of injection for softwood pulp suspension with Cm = 0.5% and for water (w), Dr = 0.05, N = 400 rpm with various mainstream velocities and almost identical jet-to-pipe velocity ratios of ~6.  271  100 R 4.05 8.26 12.4 12.4 (w)  M' (%)  10 1 0.1  Impeller 0  5  10  x/D  15  20  25  Figure B.3: Modified mixing index as a function of dimensionless distance  downstream for softwood pulp suspension with Cm = 0.5% and for water (w), Up = 2.0 m/s, Dr = 0.05, N = 400 rpm and various jet-to-pipe velocity ratios. (b)  100  M' (%)  10  N (rpm)  10  R  NA 3.63 402 3.99 603 3.86 803 3.88 803 (w) 3.88  1 Impeller  0.1 0  5  10  x/D  15  20  100  M' (%)  (a)  R  NA 6.57 209 6.12 403 6.26 601 6.20 803 6.13 805 (w) 6.0  1 0.1  0.01  25  N (rpm)  Impeller 0  5  10  x/D  15  20  25  Figure B.4: Modified mixing index as a function of dimensionless distance  downstream for softwood pulp suspension with Cm = 0.5% and for water (w), Dr = 0.05, Up = 1.0 m/s with various impeller speeds for: (a) R ≈ 4; (b) R ≈ 6. (a) U (m/s) R (b) 100  4.13 4.12 4.13 4.12  Up(m/s)  1 Impeller 0  5  10  x/D  15  20  25  R  1.0 6.21 2.0 6.13 2.0 (w) 6.20  10  M' (%)  M' (%)  10  0.1  100  p  1.0 2.0 3.0 3.0 (w)  1  0.1  Impeller 0  5  10  x/D  15  20  25  Figure B.5: Modified mixing index as a function of dimensionless distance  downstream for softwood pulp suspension with Cm = 1.0% and for water (w), Dr = 0.05, N = 400 rpm with various mainstream velocities for: (a) R ≈ 4; (b) R ≈ 6. 272  (a) 100  (b) 100  R 2.07 4.13 6.13 6.36 (w)  10  M' (%)  M' (%)  10  1  0.1  1  Impeller 0  5  10  15  x/D  20  R 4.13 6.21 8.25 12.3 12.3 (w)  0.1  25  Impeller 0  5  10  x/D  15  20  25  (c) 100  M' (%)  10  R 8.20 12.3 16.2 24.5 24.6 (w)  1 0.1 0.01  Impeller 0  5  10  x/D  15  20  25  Figure B.6: Modified mixing index as a function of dimensionless distance  downstream for softwood pulp suspension with Cm = 1.0% and for water (w), Dr = 0.05, N = 400 rpm and various jet-to-pipe velocity ratios at: (a) Up = 2.0 m/s; (b) Up = 1.0 m/s; (c) Up = 0.5 m/s.  (a)  N (rpm)  (b) 100 10  M' (%)  10  M' (%)  R  NA 3.99 404 4.13 604 4.13 803 4.13 803 (w) 4.25  100  1  N (rpm)  1 0.1  0.1  Impeller 0  5  0.01  10  x/D  15  20  25  R  NA 6.57 205 6.24 403 6.21 602 6.10 803 6.25 805 (w) 6.0  Impeller 0  5  10  x/D  15  20  25  Figure B.7: Modified mixing index as a function of dimensionless distance  downstream for softwood pulp suspension with Cm = 1.0% and for water (w), Dr = 0.05, Up = 1.0 m/s with various impeller speeds for: (a) R ≈ 4; (b) R ≈ 6. 273  (b)  (a) 100  R N (rpm) 2.07 410 4.13 420 6.20 423 6.20 (w) 403  10  M' (%)  M' (%)  10  R N (rpm) 4.09 410 6.14 402 8.05 423 12.3 430 12.3 (w) 400  100  1  1  Impeller 0.1  0  5  Impeller 10  x/D  15  20  0.1  25  0  5  10  15  x/D  20  25  Figure B.8: Modified mixing index as a function of dimensionless distance  downstream for softwood pulp with Cm = 2.0% and for water (w), Dr = 0.05, N ≈ 400 rpm and various velocity ratios for: (a) Up = 2.0 m/s; (b) Up = 1.0 m/s. 100  R  NA 410 608 804  3.99 4.09 4.09 4.14  10  1  (b)100  N (rpm)  M' (%)  M' (%)  (a)  5  R  418 610 810  24.1 24.3 24.8  10  Impeller 0  N (rpm)  Impeller  10  x/D  15  20  1  25  0  5  10  15  x/D  20  25  Figure B.9: Modified mixing index as a function of dimensionless distance  downstream for softwood pulp suspension with Cm = 2.0%, Dr = 0.05 and various impeller speeds for (a) Up = 1.0 m/s and R ≈ 4; (b) Up = 0.5 m/s and R ≈ 24. (a)  Up(m/s)  100  1.0 4.27 2.0 4.13 3.0 4.13 3.0 (w) 4.12  N (rpm) 411 400 443 440  1  0.1  (b)  5  R  0.5 8.0 1.0 8.0 1.0 (w) 8.0  N (rpm) 405 440 403  1 Impeller  Impeller 0  Up (m/s)  100  10  M' (%)  10  M' (%)  R  10  x/D  15  20  25  0.1  0  5  10  x/D  15  20  25  Figure B.10: Modified mixing index as a function of dimensionless distance  downstream for softwood pulp with Cm = 3.0% and for water (w), Dr = 0.05, almost identical impeller speeds and various mainstream velocities at: (a) R ≈ 4; (b) R ≈ 8. 274  (b)  (a) 100  R 2.07 4.13 6.20 6.20 (w)  1  R N (rpm) 4.27 411 6.14 405 8.18 433 12.3 430 12.3 (w) 400  100  403 443 440 403  10  M' (%)  10  M' (%)  N  1  Impeller 0.1  0  5  10  x/D  M' (%)  (c)  15  20  0.1  25  Impeller 0  5 R  100  10  x/D  15  20  25  N (rpm)  8.54 12.4 16.5 24.9  405 404 430 447  10  Impeller 1  0  5  10  x/D  15  20  25  Figure B.11: Modified mixing index as a function of dimensionless distance  downstream for softwood pulp suspension with Cm = 3.0% and for water (w), Dr = 0.05, almost constant impeller rotation speed and various jet velocities at: (a) Up = 2.0 m/s; (b) Up = 1.0 m/s; (c) Up = 0.5 m/s.  (a)  N (rpm)  100  100  M' (%)  M' (%)  10  1 Impeller 0.1  R  NA 4.09 411 4.27 601 4.24 808 4.24 400 (w) 4.03  0  5  10  x/D  15  20  25  N (rpm)  R  NA 0 (//) 0( ) 405 603 805  8.10 8.01 7.96 8.54 8.15 8.19  10  1  Impeller 0  5  10  x/D  15  20  25  Figure B.12: Modified mixing index as a function of dimensionless distance  downstream for softwood pulp with Cm = 3.0% and for water (w), Dr = 0.05 and various impeller speeds for: (a) Up = 1.0 m/s and R ≈ 4; (b) Up = 0.5 m/s and R ≈ 8. 275  Up (m/s)  100  10  M' (%)  R  N (rpm)  3.0 4.11 1.0 4.13 2.0 4.12 3.0 4.17 3.0 (w) 4.12  NA 406 410 409 404  1 Impeller 0.1  0  5  10  x/D  15  20  25  Figure B.13: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension with Cm = 0.5% and for water (w), Dr = 0.05, N ≈ 400 rpm and R ≈ 4, with various mainstream velocities. (b)  (a) 100 R  2.00 408 4.12 410 6.13 406 6.20 (w) 403  1 Impeller 0.1  0  5  R  1 0.1  10  x/D  15  20  25  N (rpm)  7.96 12.3 16.4 24.8 25.1 (w)  10  M' (%)  M' (%)  10  100  N (rpm)  403 406 407 407 402  Impeller 0  5  10  x/D  15  20  25  Figure B.14: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension with Cm = 0.5% and for water (w), Dr = 0.05, N ≈ 400 rpm and various velocity ratios at: (a) Up = 2.0 m/s; (b) Up = 0.5 m/s. N (rpm)  10  M' (%)  R  0 (//) 6.25 0 ( ) 6.11 220 6.13 408 5.98 601 6.0 803 6.11 805 (w) 6.0  100  1 0.1 0.01  Impeller 0  5  10  x/D  15  20  25  Figure B.15: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension with Cm = 0.5% and for water (w), Dr = 0.05, Up = 1.0 m/s and R ≈ 6 with various impeller speeds. 276  100 100  Up(m/s) 1.0 2.0 3.0 3.0 (w)  Up(m/s)  4.13 4.12 4.0 4.12  R  1.0 6.25 2.0 6.19 2.0 (w) 6.20  10  M' (%)  M' (%)  10  R  1  1  Impeller  Impeller  0.1 0  5  10  15  20  0.1  25  0  5  10  15  x/D  20  25  x/D  Figure B.16: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension with Cm = 1.0% and for water (w), Dr = 0.05, N = 400 rpm and various mainstream velocities at: (a) R ≈ 4; (b) R ≈ 6.  (b) 100  (a) 100  R  R 2.07 4.12 6.19 6.20 (w)  10  M' (%)  M' (%)  10  1  1 Impeller  Impeller 0.1  0  5  N  4.13 417 6.25 404 8.22 408 12.3 405 12.3 (w) 400  10  x/D  (c)  15  20  0.1  25  100  5  10  R  N  x/D  15  20  25  8.0 407 12.5 406 16.4 410 24.5 404 24.6 (w) 402  10  M' (%)  0  1 0.1 Impeller 0.01  0  5  10  x/D  15  20  25  Figure B.17: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension with Cm = 1.0% and for water (w), Dr = 0.05, N ≈ 400 rpm and various velocity ratios at: (a) Up = 2.0 m/s; (b) Up = 1.0 m/s; (c) Up = 0.5 m/s.  277  (a)  (b) 100  100  M' (%)  R  NA 4.25 0 (//) 4.28 0 ( ) 4.01 417 4.13 606 4.13 802 4.02 803 (w) 4.25  1  0  5  10  x/D  15  20  Impeller 0  25  5  R  NA 6.18 0 (//) 6.26 0( ) 6.25 404 6.25 601 6.26 803 6.27 805 (w) 6.0  1 0.1  Impeller  0.1  N (rpm)  10  M' (%)  N (rpm)  10  10  x/D  15  20  25  (c) 100 N (rpm) R  M' (%)  10  NA 0 (//) 0( ) 407 601 802  1  8.45 7.95 8.0 7.99 7.95 8.02  Impeller 0.1  0  5  10  x/D  15  20  25  Figure B.18: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension with Cm = 1.0% and for water (w), Dr = 0.05, Up = 1.0 m/s and various impeller speeds at: (a) R ≈ 4; (b) R ≈ 6; (c) R ≈ 8. Up(m/s)  100  1.0 2.0 3.0 3.0 (w)  M' (%)  10  R  N (rpm)  4.09 3.94 4.01 4.12  437 437 430 404  1  0.1  (b) 100 Up(m/s)  5  N (rpm) 423 430 403  1 Impeller  Impeller 0  R  1.0 6.26 2.0 6.19 2.0 (w) 6.20  10  M' (%)  (a)  10  x/D  15  20  25  0.1  0  5  10  x/D  15  20  25  Figure B.19: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension with Cm = 2.0% and for water (w), Dr = 0.05, almost identical impeller rotation speeds and various mainstream velocities at: (a) R ≈ 4; (b) R ≈ 6. 278  (a) 100  (b) 2.06 3.94 6.19 6.20 (w)  10  M' (%)  N (rpm) 447 430 430 400  R N (rpm) 4.09 437 6.26 423 8.26 443 12.3 443 12.3 (w) 400  100  10  M' (%)  R  1  1  Impeller 0.1  0  5  Impeller  10  x/D  15  20  0.1  25  0  5  10  x/D  15  20  25  (c) 100  M' (%)  10  R N (rpm) 8.0 430 12.5 417 16.4 437 24.5 440 24.6 (w) 402  1 0.1 Impeller 0.01  0  5  10  x/D  15  20  25  Figure B.20: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension with Cm = 2.0% and for water (w), Dr = 0.05, almost identical impeller rotation speeds and various velocity ratios at: (a) Up = 2.0 m/s; (b) Up = 1.0 m/s; (c) Up = 0.5 m/s. (a)  100  100  M' (%)  1 0.1  Impeller 0  5  R  NA 2.06 0 (//) 2.06 0 ( ) 2.06 447 2.06 600 2.06 800 2.06 801 (w) 2.08  10  x/D  15  20  25  10  M' (%)  N (rpm)  10  N (rpm)  R  NA 4.09 437 4.09 610 4.09 830 4.09 803 (w) 4.25  1 0.1  Impeller 0  5  10  x/D  15  20  25  Figure B.21: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension with Cm = 2.0% and for water (w), Dr = 0.05 and various impeller speeds at: (a) Up = 2.0 m/s and R ≈ 2; (b) Up = 1.0 m/s R ≈ 4. 279  (b)  100  100 N (rpm)  M' (%)  10  R  NA 6.26 0 (//) 6.23 0 ( ) 6.16 423 6.26 613 6.26 807 6.30 805 (w) 6.0  1 0.1  N (rpm) R  10  M' (%)  (a)  NA 0 (//) 0( ) 407 601 802  1  Impeller 0.01  0  5  10  x/D  15  20  0.1  25  8.18 8.10 8.10 8.10 8.10 8.10  Impeller 0  5  10  x/D  15  20  25  Figure B.22: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension with Cm = 2.0% and for water (w), Dr = 0.05 and various impeller speeds at: (a) Up = 1.0 m/s and R ≈ 6; (b) Up = 0.5 m/s R ≈ 8.  (a)  Up(m/s)  100  1.0 2.0 3.0 3.0 (w)  N (rpm)  4.05 4.13 4.12 4.12  433 445 443 404  (b) 100  1  Up(m/s)  0  5  N (rpm) 434 445 403  1  Impeller 0.1  R  1.0 6.11 2.0 6.19 2.0 (w) 6.20  10  M' (%)  M' (%)  10  R  Impeller  10  x/D  15  20  25  0.1  0  5  10  x/D  15  20  25  Figure B.23: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension with Cm = 3.0% and for water (w), Dr = 0.05, almost identical impeller rotation speeds and various mainstream velocities at: (a) R ≈ 4; (b) R ≈ 6.  280  (b)  (a) 100  R 2.07 4.13 6.19 6.20 (w)  445 445 445 400  1  1 Impeller  Impeller 0.1  0  5  N  4.05 433 6.11 434 8.47 440 12.4 432 12.3 (w) 400  10  M' (%)  M' (%)  10  R  100  N (rpm)  10  x/D  15  20  0.1  25  0  5  M' (%)  (c) 100  10  x/D  15  20  25  R N (rpm) 8.49 435 11.8 420 16.7 430 24.5 435 24.6 (w) 402  10  Impeller 1  0  5  10  x/D  15  20  25  Figure B.24: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension with Cm = 3.0% and for water (w), Dr = 0.05, almost identical impeller rotation speeds and various jet velocities at: (a) Up = 2.0 m/s; (b) Up = 1.0 m/s; (c) Up = 0.5 m/s. 100  M' (%)  10  N (rpm)  1 0.1 Impeller 0.01  0  5  R  NA 6.14 0 (//) 5.95 0 ( ) 6.0 434 6.11 604 5.89 805 5.89 805 (w) 6.0  10  15  20  25  x/D  Figure B.25: Modified mixing index as a function of dimensionless distance  downstream for hardwood pulp suspension with Cm = 3.0% and for water (w), Dr = 0.05, Up = 1.0 m/s and R ≈ 6 with various impeller speeds. 281  Appendix C: Supplementary Data for Chapter 6 This appendix includes additional experimental results for gas dispersion in water and softwood pulp suspension horizontal flow.  (b) 100  (a) 100 S EB/S EB/S  60 Mg (%)  S  S  EB/S EB B B  20  EB Usg(m/s) 0.11 0.22 0.33 0.44  B  0  B/EB B/EB B/EB  EB  B  40  10  S  5  10  x/D  20  10 5  25  B  0  DB DB  B  5  DB  B B  DB  B  B  0.11 0.22 0.33 0.44  DB  5  10  x/D  Usg(m/s)  15  20  25  Usg(m/s)  B  10  \  B  DB  100  Mg (%)  (c)  15  B  Mg (%)  80  0  DB DB B B  B B  DB DB DB B  5  10  15  B B B  20  0.11 0.22 0.33 0.44  B B B  25  x/D  Figure C.1: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for air-water flow for various superficial gas velocities at: (a) Usl = 1.0 m/s; (b) Usl = 3.0 m/s; (c) Usl = 4.0 m/s. Letters identify flow regimes as in  Figure 6.6.  282  (b)  (a)  Mg (%)  B B  10 5  S  S  S  EB/S EB/S S  S  S  EB/S EB/S  B DB DB DB  0  S  DB DB  B/EB  B B  10  x/D  Usl(m/s) 0.5 1.0 2.0 3.0 4.0 5.0  B/EB B/EB B  B  B  DB DB  5  EB EB EB  B/EB B/EB  B  DB DB  15  100 80 60  S  20  Mg (%)  100  EB/S EB/S  40 20 10  25  S S S S S S S S B/EB B/EB EB EB B B B/EB B B/EB B/EB DB DB B B B B B DB B B DB DB DB DB EB/S EB/S S  0  5  10  x/D  15  20  Usl(m/s) 0.5 1.0 2.0 3.0 4.0 5.0  25  (c)  Mg (%)  100 80 60  EB/S EB/S S  S S S EB/S S EB/S B EB B/EB B/EB EB B B/EB DB B/EB B/EB  40  B  20  DB  DB  0  B  DB  DB  10  S  S  S  S  B  B  DB DB DB  5  10  x/D  B  B  B  B  15  20  Usl(m/s) 0.5 1.0 2.0 3.0 4.0 5.0  25  Figure C.2: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for air-water flow for various superficial liquid velocities at: (a) Usg = 0.22 m/s; (b) Usg = 0.33 m/s; (c) Usg = 0.44 m/s. Letters identify flow regimes  as in Figure 6.6.   100 EB/SEB/S  S  S  S  S  Mg (%)  EB/S EB/S EB/S EB EB  B B  10  B  DB DB DB  4  0  5  B DB  EB B/EB B/EB  B  DB DB  x/D  S EB EB B/EB  B B/EB B B B  DB DB DB  10  Usl(m/s) T ( C)  15  DB DB  20  0.5 0.5 1.0 1.0 2.0 2.0 3.0 3.0 4.0 4.0 5.0 5.0  15 25 15 25 15 25 15 25 15 25 15 25  25  Figure C.3: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for air-water flow at Usg = 0.11 m/s for various superficial liquid velocities and mainstream temperatures. 283  (a) 0.36  (b) 0.16  0.18 0.12  Local gas holdup  Local gas holdup  2.41 5.69 8.97 12.2 15.5 18.8 22.1  0.24  0.06 0.00  x/D  x/D  0.30  0  20  40  60  0.08 0.04 0.00  80  2.41 5.69 8.97 12.2 15.5 18.8 22.1  0.12  Distance from top of pipe (mm)  0  20  Local gas holdup  (c)0.06  60  80  x/D 2.41 5.69 8.97 12.2 15.5 18.8 22.1  0.04  0.02  0.00  40  Distance from top of pipe (mm)  0  20  40  60  80  Distance from top of pipe (mm)  Figure C.4: Vertical gas holdup profiles as a function of dimensionless distance  downstream for air-water flow at Usg = 0.11 m/s for: (a) Usl = 2.0 m/s; (b) Usl = 3.0 m/s; (c) Usl = 4.0 m/s.  (a)  (b) 100  100 90  80 60  Stratified flow  80  Usg(m/s) 0.11 0.22 0.33 0.44  50 40  Mg (%)  Mg (%)  70 60  5  10  15  20  25  S  S  S EB  B  40  20 10  0  EB/S EB/S S  EB/S EB EB Usg(m/s) 0.11 0.22 0.33 0.44  B  0  x/D  5  10  15  20  25  x/D  Figure C.5: Gas mixing index as a function of dimensionless distance downstream  for Cm = 0.5% and various superficial gas velocities at: (a) Usl = 0.5 m/s; (b) Usl = 1.0 m/s. Letters identify flow regimes as in Figure 6.6. 284  (b) 100  (a) 100 80 B B DB  B/EB B/EB EB  Usg(m/s)  B  20 B  0  0.11 0.22 0.33 0.44  DB  5  EB  10  x/D  15  20  Usg(m/s)  B  Mg (%)  DB DB B  10 5  0  B B  DB DB DB  5  10  B B B  15  x/D  B B B  20  B  10 5  25  100  B  Mg (%)  40  10  (c)  B/EB B/EB EB EB  0.11 0.22 0.33 0.44  B B B  25  B  0  DB DB DB  5  (d) 100  Mg (%)  Mg (%)  60  B B B  DB  B/EB B/EB  Usg(m/s)  B  0.11 0.22 0.33 0.44  DB  10  x/D  15  20  25  Usg(m/s)  B  10 4  DB  B/EB  DB  0.11 0.22 0.28 0.33 0.44  DB DB DB  DB  B DB DB  0  B  5  DB  10  DB  15  DB  20  B  B B  25  x/D  Figure C.6: Gas mixing index as a function of dimensionless distance downstream  for Cm = 0.5% and various superficial gas velocities at: (a) Usl = 2.0 m/s; (b) Usl = 3.0 m/s; (c) Usl = 4.0 m/s; (d) Usl = 5.0 m/s. Letters identify flow regimes as in Figure 6.6.  285  (b)  (a)  EB/S EB/S  EB/S EB/S EB/S EB EB  Mg (%)  B DB DB B B DB DB B DB DB DB DB DB  10 4  0  5  EB  0.5 1.0 2.0 3.0 4.0 5.0  10  x/D  15  B  B  B  5  25  DB  DB  0  DB  B  DB  B  10  0.5 1.0 2.0 3.0 4.0 5.0  B/EB B/EB B/EB  B  B  B  DB DB DB DB  5  Usl(m/s)  15  B  20  25  x/D  (d) EB/S EB/S S  40  B  S  S  S  EB/S EB/S  B DB  20 B  B/EB B/EB EB EB B  DB DB  DB  0  5  S  B/EB  B  DB DB  10  0.5 1.0 2.0 3.0 4.0 5.0  B/EB B/EB  B B B B B DB  15  x/D  Usl(m/s)  20  25  100 80 60  Mg (%)  100 80 60  Mg (%)  S S  B  DB  10  (c)  10  B  DB DB  20  S EB EB EB  S S  S  S  EB/S EB/S S B/EB B/EB  Usl(m/s)  B/EB B/EB B/EB B B/EB  B  EB/S EB/S  EB  EB  B  100  S  S  S  S  S  Mg (%)  100  EB/S EB/S S  40  B B  20  EB/S  DB DB DB  10  0  5  EB/S  B B/EB  DB DB  B B  x/D  S  B/EB EB EB  B/EB  DB DB  10  S  S  S  0.5 1.0 2.0 3.0 4.0 5.0  B/EB B/EB  B DB  15  Usl(m/s)  B B  20  B B  25  Figure C.7: Gas mixing index as a function of dimensionless distance downstream  for Cm = 0.5% and various superficial liquid velocities at: (a) Usg = 0.11 m/s; (b) Usg = 0.22 m/s; (c) Usg = 0.33 m/s; (d) Usg = 0.44 m/s. Letters identify flow regimes as in Figure 6.6.  286  (b) 100  100 80  Mg (%)  60  EB EB/S EB  40  20  S  EB/S S  0  EB  EB  EB EB  80  EB  EB  EB EB  40  Usg(m/s)  EB  5  0.11 0.22 0.33 0.44  10  15  x/D  20  EB/S  Usg(m/s) 0.11 0.22 0.33 0.44  B  0  EB  EB EB  5  10  15  20  25  (d) 100 80 B  40  B  EB EB  B/EB EB EB  DB  EB  B/EB  20  0.11 0.22 0.33 0.44  DB  0  5  (e) 100  EB  Usg(m/s)  B B  60  10  x/D  15  20  25  Mg (%)  DB DB  B  B B  10  B DB  0  5  DB  B B B  B  15  10  20  5  B  25  x/D  5  Usg(m/s) 0.11 0.22 0.33 0.44  DB  10  x/D  15  20  25  Usg(m/s)  DB  0.11 0.22 0.33 0.44  DB DB  0  B  B  B  DB  10  DB  100  B  B B  0  EB B/EB  DB  20  B B  B/EB  B  B DB  10  (f)  B  40  Mg (%)  B  B/EB  B  Usg(m/s) 0.11 0.22 0.33 0.44  Mg (%)  60  Mg (%)  EB/S EB/S  S  S  S  x/D  80  5  EB/S S  B  20  10  25  (c) 100  10  EB/S  60  EB  Mg (%)  (a)  DB  DB DB  5  10  DB  15  B  B B  20  B  B B  25  x/D  Figure C.8: Gas mixing index as a function of dimensionless distance downstream  for Cm = 1.0% and various superficial gas velocities at: (a) Usl = 0.5 m/s; (b) Usl = 1.0 m/s; (c) Usl = 2.0 m/s; (d) Usl = 3.0 m/s; (e) Usl = 4.0 m/s; (f) Usl = 5.0 m/s. Letters identify flow regimes as in Figure 6.6.  287  (a)100  (b) 100  Usg(m/s) 0.11 0.22 0.33 0.44  80  Mg (%)  60  EB  EB  EB EB  40 EB/S EB  20  0  5  EB  EB EB  EB  10  15  x/D  EB  20  EB  EB  B  Usg(m/s) 0.11 0.22 0.33 0.44  B DB  0  5  (e) 100  x/D  15  20  0.11 0.22 0.33 0.44  60 40  B B  20  B B  B B  0  5  B  B  5  B  10  15  20  25  10  0.11 0.22 0.33 0.44  B  x/D  15  20  EB  25  EB  B/EB B/EB  B/EB  DB  B DB  DB  5  10  0  B/EB  x/D  15  B/EB  20  25  Usg(m/s)  100  0.11 0.22 0.33 0.44  B B  B  B  B B  0  5  x/D  B  10  B  B  B  B  5  Usg(m/s)  EB  DB  10  B  EB EB  Usg(m/s)  20 10  (f)  B/EB  EB 0.11 0.22 0.33 0.44  0  40  B/EB B/EB B  EB  B  25  Usg(m/s)  80  Mg (%)  10  EB  B  60  EB  EB  DB  20  EB EB  Mg (%)  EB  EB  Mg (%)  Mg (%)  B  B  40  10  20  80  60  EB EB EB EB  EB/S  (d) 100  80  S  B  40  10  25  (c) 100  10  EB/S  60  Mg (%)  EB/S EB EB  80  B B  B  B  x/D  B B  15  20  25  Figure C.9: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 2.0% and various superficial gas velocities at: (a) Usl = 0.5 m/s; (b) Usl = 1.0 m/s; (c) Usl = 2.0 m/s; (d) Usl = 3.0 m/s; (e) Usl = 4.0 m/s; (f) Usl = 5.0 m/s. Letters identify flow regimes as in Figure 6.6.  288  (a)  Usg(m/s)  100 80  Mg (%)  60  EB EB EB  EB/S  EB  80 60  40 EB/S EB EB EB  20  0  5  10  EB  15  EB  (b) 100  EB  20  EB EB EB EB EB EB  B  40  Mg (%)  EB EB  0.11 0.22 0.33 0.44  B  20 10  25  0  5  (d) 100  10  20  B  0  EB EB EB B/EB  B  5  60  EB  Usg(m/s)  EB  B  10  0.11 0.22 0.33 0.44  15  20  Mg (%)  Mg (%)  B  40  10  EB EB EB  40  B  B  20  10  25  B  0.11 0.22 0.33 0.44  B  20  25  (e) 100  B/EB B/EB B/EB B  B B  0  5  10  x/D  B  B  x/D  15  B  20  B  25  Usg(m/s)  80 60  Mg (%)  15  Usg(m/s)  80  80 EB  Usg(m/s)  x/D  (c) 100 B  EB 0.11 0.22 0.33 0.44  x/D  60  EB EB EB  EB EB  B  40  B  20  B  B  B  B  B  B  B B  0  0.11 0.22 0.33 0.44  5  B  10  B  15  B  20  B  25  x/D  Figure C.10: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 3.0% and various superficial gas velocities at: (a) Usl = 0.5 m/s; (b) Usl = 1.0 m/s; (c) Usl = 2.0 m/s; (d) Usl = 4.0 m/s; (e) Usl = 5.0 m/s. Letters identify flow regimes as in Figure 6.6.  289  Appendix D: Supplementary Data for Chapter 7 This appendix includes additional experimental results for mixing gas into water and into softwood pulp suspensions using an in-line mechanical mixer.  (b) 100  100 90 80 70 EB/S  S  S  S  S  S  S  Usg(m/s)  60  0.055 0.11 0.22 0.33 0.44  50 40 Impeller 30  DB/EB  0  5  10  x/D  15  20  DB/EB DB  DB  Mg (%)  Mg (%)  (a)  DB  10  DB  DB DB DB DB  0.11 0.22 0.33 0.44  DB  DB  Impeller 5 0 5 10  25  B/EB B/EBB/EB B/EB B B B B B Usg (m/s)  x/D  15  20  25  Figure D.1: Gas mixing index and flow regimes as a function of dimensionless  distance downstream in air-water flow for N = 400 rpm and various superficial gas velocities at: (a) Usl = 0.5 m/s; (b) Usl = 2.0 m/s. Letters identify flow regimes as in Figure 6.6. 100  Mg (%)  B  B  B B  B DB 10  EB  EB B/EB B/EB  N (rpm) NA 201 401 600 803  DB DB  DB DB Impeller 5 0 5 10  15  20  25  x/D Figure D.2: Gas mixing index and flow regimes as a function of dimensionless  distance downstream in air-water flow for Usl = 2.0 m/s, Usg = 0.11 m/s and various impeller speeds. Letters identify flow regimes as in Figure 6.6. 290  EB/S S  S  S  S  Mg (%)  DB  B  5 0  DB B DB DB DB DB DB DB  Usg (m/s) 0.5 1.0 2.0 3.0  Impeller 5  10  x/D  DB/EB B/EB DB  15  20  DB DB  DB DB DB  10  B  S  S  S  S  B/EB DB/EB  DB  10  S  S  EB/S  B/EB  B  B  (b)100  S  S  Mg (%)  (a)100  B/EB  B  5 0  5  10  x/D  15  Usl (m/s) 0.5 1.0 2.0 3.0  Impeller  25  B  20  25  Mg (%)  (c) 100  S 80 EB/S S S S S S 60 B/EB B/EB B/EB DB/EB DB/EB 40 B DB B DB DB Usl(m/s) B B 0.5 20 DB DB  10  1.0 2.0 3.0  Impeller 0  5  10  x/D  15  20  25  Figure D.3: Gas mixing index and flow regimes as a function of dimensionless  distance downstream in air-water flow for N = 400 rpm and various superficial liquid velocities at (a) Usg = 0.22 m/s; (a) Usg = 0.33 m/s; (a) Usg = 0.44 m/s. Letters identify flow regimes as in Figure 6.6.   Average gas holdup  0.18 0.15 0.12 0.09  Usl (m/s) T ( C) 2.0 2.0 2.0 3.0 3.0 3.0  14.9 19.8 24.2 15.0 20.1 24.4  0.1  0.2  0.06 0.03 0.00 0.0  0.3  0.4  0.5  Usg (m/s)  Figure D.4: Average gas holdup as a function superficial gas velocity for various  superficial liquid velocities and mainstream temperatures in air-water flow. 291  100 80 60  B/EB B/EB B/EB DB/EB DB/EB DB DB  Mg (%)  40    T ( C)  20  15.0 19.8 24.2  Impeller 10  0  5  10  x/D  15  20  25  Figure D.5: Gas mixing index as a function of dimensionless distance downstream  in air-water flow for Usl = 2.0 m/s, Usg = 0.44 m/s, N = 400 rpm and various mainstream temperatures. Letters identify flow regimes as in Figure 6.6. 100  EB/S  S  S  S  S  S S EB/S EB/S o  Usl(m/s) T ( C)  Mg (%)  EB/S EB/S EB/S B B B DB  DB DB DB DB DB DB DB  10 5  0  Impeller 5  10  DB  x/D  DB  B  DB  15  0.5 0.5 1.0 1.0 2.0 2.0 3.0 3.0  B  20  15.2 24.4 15.0 24.2 14.7 24.2 14.7 24.3  25  Figure D.6: Gas mixing index and flow regimes as a function of dimensionless  distance downstream in air-water flow for Usg = 0.11 m/s, N = 400 rpm and various superficial liquid velocities and mainstream temperatures. 100 90 80 S 70 EB/S  (b) 100 EB/S  S  S  S  S  S  60  Usg(m/s)  50  0.055 0.11 0.22 0.33 0.44  40 Impeller  30  0  5  10  x/D  15  20  EB/S  S  S  S  S  S  EB/S EB/S EB/S EB/S EB/S EB/S EB/S EB/S EB/S Usg (m/s)  B EB/S DB  Mg (%)  Mg (%)  (a)  B  10  0.055 0.11 0.22 0.33 0.44  DB  Impeller  25  0  5  10  x/D  15  20  25  Figure D.7: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 0.5%, N = 400 rpm and various superficial gas velocities at: (a) Usl = 0.5 m/s; (b) Usl = 1.0 m/s. 292  (b) 100  (a) 100  Mg (%)  DB DB  DB  B DB  DB  DB  DB  10  B/EB B/EB  B/EB  Usg (m/s)  DB  0.11 0.22 0.33 0.44  DB DB  Impeller 0 5 10  x/D  DB  B  B  B  15  0.11 0.22 0.33 0.44  B/EB  Mg (%)  B/EB  Usg (m/s)  DB DB  10  20  DB DB Impeller  0  25  B  B DB DB  DB  5  10  B  B  DB DB  B B  B B  DB  DB  DB  DB  x/D  B  B  15  20  25  Figure D.8: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 0.5%, N = 400 rpm and various superficial gas velocities at: (a) Usl = 2.0 m/s; (b) Usl = 3.0 m/s. EB/S S  S  S  S  S  (b) 100  S  EB/SEB/S EB/S EB/S B B EB/S B Usl(m/s) B DB 0.5 DB 1.0 B DB 2.0 DB DB 3.0 DB DB 10 DB DB DB DB 5 Impeller 0 5 10 15 20 25  DB  S  S  S  B/EB B/EB DB  DB  20  10  0  5  S  B Usl (m/s)  B DB  10  0.5 1.0 2.0 3.0  DB  x/D  15  S  B/EB B/EB B  DB  DB Impeller  DB  20  25  B  B  Usg (m/s)  DB  DB  0.5 1.0 2.0 3.0  Impeller  5 0  5  10  x/D  15  20  (d) 100  Mg (%)  Mg (%)  100 80 EB/S S 60 EB/S B/EB 40  DB  DB DB DB  10  B/EB B/EB  B  DB  S  S  S  S  EB/S  x/D  (c)  S  EB/S S  Mg (%)  Mg (%)  (a) 100  80 EB/S S 60 B/EB 40 DB DB  20  S  B/EB DB  DB DB  S  S  B/EB  B  B  S B/EB  0  5  10  x/D  S B/EB B  B  Usl(m/s) 0.5 1.0 2.0 3.0  Impeller 10  25  15  20  25  Figure D.9: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 0.5%, N = 400 rpm and various superficial liquid/pulp velocities at: (a) Usg = 0.11 m/s; (b) Usg = 0.22 m/s; (c) Usg = 0.33 m/s; (d) Usg = 0.44 m/s. Letters identify flow regimes as in Figure 6.6. 293  (b) 100  (a) 100  60  S  N (rpm)  EB/S  50 EB/S  NA 0 (//) 0( ) 404 602 803  40 Impeller 30  0  5  S  S  S  S  Mg (%)  Mg (%)  90 80 EB/S S 70  10  x/D  15  20  0  EB  NA 405 600 801  DB DB DB Impeller  5  B  N (rpm)  DB  10  5  10  x/D  15  20  25  N (rpm) NA 0 (//) 0( ) 408 603 803  Mg (%)  B DB  25  (c) 100  B/EB B/EB EB B B B  B  10  DB  DB  DB DB  DB  B  B B/EB B DB  10  B  DB  B DB  DB  5 Impeller 0 5  B/EB  x/D  15  DB  20  25  Figure D.10: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 0.5%, Usg = 0.11 m/s and various impeller rotation speeds at: (a) Usl = 0.5 m/s; (b) Usl = 2.0 m/s; (c) Usl = 3.0 m/s. Letters identify flow regimes as in Figure 6.6. EB/S S  Mg (%)  B/EB  S  S  (b) 100  S  S  B/EB B/EB  EB EB EB EB EB  Mg (%)  (a) 100  Usg (m/s) N (rpm)  10 5  0.055 0.11 0.22 0.33 0.44  DB  Impeller 0  5  10  x/D  15  440 423 437 440 440  20  10  25  5  B  0.11 0.22 0.33 0.44  DB DB  Impeller  0  5  EB EB  EB  B/EB B/EB B/EB B/EB B Usg (m/s) N (rpm)  B DB  EB  10  x/D  15  20  437 440 440 437  25  Figure D.11: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 1.0%, virtually identical impeller speeds and various superficial gas velocities at: (a) Usl = 1.0 m/s; (b) Usl = 2.0 m/s. 294  S  EB/S  S  EB/S  S  Mg (%)  B  Usl(m/s) N (rpm)  B  DB DB Impeller  0  5  B  B  DB  5  80 EB/S S S 60 EB/S  EB EB B/EB EB  B/EB  10  (b)100  S  S  S  10  0.5 1.0 2.0 3.0  DB  x/D  15  433 437 440 428  20  Mg (%)  (a)100  20  EB  B  DB  10  25  B/EB  40  S  DB  B  EB EB EB B/EB B/EB B Usl(m/s) N (rpm)  DB  0.5 1.0 2.0 3.0  Impeller 0  5  10  S  S  S  x/D  15  440 440 440 428  20  25  Mg (%)  (c) 100  S S 80 EB/S S S S S EB EB 60 EB/S EB/S EB EB B/EB B/EB B/EB 40 B/EB B/EB B Usl(m/s) N (rpm) B DB 0.5 432 20 1.0 2.0 3.0  Impeller 10  0  5  10  x/D  15  20  440 437 432  25  Figure D.12: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 1.0%, virtually identical impeller speeds and various superficial liquid/pulp velocities at (a) Usg = 0.22 m/s; (b) Usg = 0.33 m/s; (c) Usg = 0.44 m/s. Letters identify flow regimes as in Figure 6.6. 100 90 80 EB/S S 70 60 EB/S 50 EB  40  (b) S  S  S S  S  S  EB EB  S EB N (rpm) NA 0 (//) 0( ) 437 602 804  EB EB  Impeller 30  0  5  10  x/D  15  20  25  Mg (%)  Mg (%)  (a)  100 80 S S S S S EB/S EB/S 60 S S S S S S EB EB/S EB/S 40 EB EB EB/S EB/S 20 10  N (rpm) NA 0 (//) 0( ) 423 600 807  B  Impeller 0  5  10  x/D  15  20  25  Figure D.13: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 1.0%, Usg = 0.11 m/s and various impeller speeds at: (a) Usl = 0.5 m/s; (b) Usl = 1.0 m/s. Letters identify flow regimes as in Figure 6.6. 295  (b)  100  10 5  DB  NA 437 603 801  DB DB  Impeller 0 5 10  x/D  15  20  N (rpm)  100  Mg (%)  B/EB EB B/EB B/EB EB B EB B/EB B N (rpm) DB B  Mg (%)  (a)  NA 0 (//) 0( ) 433 607 802  B  DB  DB  B/EBB/EB B B  B  B  B  DB  B  DB  10  B DB DB DB DB 5 Impeller DB 0 5 10 15 20  25  25  x/D  Figure D.14: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 1.0%, Usg = 0.11 m/s and various impeller speeds at: (a) Usl = 2.0 m/s; (b) Usl = 3.0 m/s. Letters identify flow regimes as in Figure 6.6. EB/S  Mg (%)  EB  DB DB  10 5  (b) 100  S  B  S  S EB  EB  B/EB  EB  EB EB/S EB EB EB EB B/EB B/EB Usl(m/s) N (rpm)  B  0.5 1.0 2.0 3.0  DB  DB  DB  Impeller 0 5  10  x/D  15  20  421 401 416 427  EB B  20 10  S  S  S  S  S  EB EB EB B/EB B/EB  DB DB  B  x/D  EB  B/EB  Usg (m/s) N (rpm)  B/EB  Impeller 0 5 10  EB  EB  15  0.5 1.0 2.0 3.0  418 412 415 433  20  25  (d) 100 S  EB/S  60  S  S  S  EB  80 S  EB EB EB/S EB EB EB B/EB EB B/EB DB B/EB Usl(m/s) N (rpm) DB B 0.5 411  40  20  1.0 2.0 3.0  Impeller 0  5  10  x/D  15  20  408 420 430  25  EB/S  S  60 EB/S  EB  Mg (%)  80  Mg (%)  40  25  (c)100  10  80 EB/S 60  S  Mg (%)  (a) 100  S  S  EB B/EB EB  40  EB B/EB  20 10  S  DB  EB EB  5  10  S EB  EB  Usl(m/s) N (rpm)  B  0.5 1.0 2.0 3.0  Impeller 0  S  x/D  15  413 417 420 440  20  25  Figure D.15: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 2.0%, virtually identical impeller speeds and various superficial liquid/pulp velocities at: (a) Usg = 0.11 m/s; (b) Usg = 0.22 m/s; (c) Usg = 0.33 m/s,; (d) Usg = 0.44 m/s. 296  (a)  (b) 100 S  S  S S  EB/S  40  B/EB EB EB  20  EB  B/EB EB EB  S S EB/S  EB  Usg (m/s) N (rpm) 0.055 0.055 0.11 0.22 0.33 0.44  EB EB EB  DB  10 0  NA 401 412 408 417 417  10  x/D  15  20  EB EB  B  10  Impeller  Impeller 5  B/EB  DB  5  25  0  5  10  x/D  EB  EB EB  B/EB  B  Mg (%)  Mg (%)  100 80 S EB/S 60  EB EB Usg (m/s) N (rpm) 0.11 0.22 0.33 0.44  416 415 420 420  20  25  15  Figure D.16: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 2.0%, almost identical impeller speeds and various superficial gas velocities at: (a) Usl = 1.0 m/s; (b) Usl = 2.0 m/s. Letters identify flow regimes as in Figure 6.6.  (b)  (a)  S EB/S EB/S  40 EB/S EB  20  N (rpm)  EB  NA 0 (//) 0( ) 421 602 801  EB EB EB EB EB  Impeller 10  0  5  100  10  x/D  15  DB  5  5  NA 416 615 802  DB  DB  Impeller 0  5  10  x/D  15  20  25  N (rpm)  B  10  N (rpm)  B  10  25  100  Mg (%)  (b)  20  EB EB EB EB B/EB  B  Mg (%)  Mg (%)  100 80 EB/S S 60  DB DB  NA 427 603 810  B  DB  B/EB B/EB B/EB B  B/EB  B/EB  B  DB Impeller 0 5 10  x/D  15  20  25  Figure D.17: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 2.0%, Usg = 0.11 m/s and various impeller speeds at (a) Usl = 0.5 m/s; (b) Usl = 2.0 m/s; (c) Usl = 3.0 m/s. 297  (b)  (a)  100 80 S 60 EB/S  EB  20  EB  Usg (m/s) N (rpm)  EB EB EB  0.055 0.11 0.22 0.33 0.44  EB  Impeller 10  0  5  10  x/D  15  20  430 427 430 430 427  S  40  Mg (%)  Mg (%)  100 80 EB/S EB EB EB EB EB EB 60 EB 40 EB  Usg (m/s) N (rpm)  20 Impeller 10  25  0  EB EB  EB  EB  0.055 0.11 0.22 0.33 0.44  410 423 433 425 432  EB EB EB EB EB EB EB EB EB EB EB EB EB EB EB  5  10  x/D  15  20  25  (c)100 80 B EB EB EB EB EB  60 EB/S  Mg (%)  40 EB EB  20  B/EB  Usg (m/s) N (rpm)  EB EB  0.11 0.22 0.33 0.44  Impeller  10 0  5  10  x/D  EB  15  418 437 440 447  20  25  Figure D.18: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 3.0%, almost identical impeller speeds and various superficial gas velocities at: (a) Usl = 0.5 m/s; (b) Usl = 1.0 m/s; (c) Usl = 2.0 m/s. Letters identify flow regimes as in Figure 6.6.  (b)  (a)  40  100  S S EB  S  S  S EB  NA 0 (//) 0( ) 423 601 809  EB EB EB EB EB EB EB  Impeller 10  0  5  10  x/D  15  20  S  S  25  S  EB/S  N (rpm)  EB EB EB  B  20  S  S  S  B/EB EB EB  B BB  Mg (%)  Mg (%)  100 80 EB/S 60  B B  S EB  EB EB  N (rpm) NA 0 (//) 0( ) 418 604 803  B/EB  10 Impeller  5 0  5  10  x/D  15  20  25  Figure D.19: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 3.0%, Usg = 0.11 m/s and various impeller speeds at: (a) Usl = 1.0 m/s; (b) Usl = 2.0 m/s. Letters identify flow regimes as in Figure 6.6. 298  100 80 S 60 EB/S 40  (b)100 EB  EB  EB EB EB EB  B  B/EB B B/EB  20  B/EB  Usl(m/s) N (rpm) 0.5 1.0 2.0 3.0  DB DB  10 Impeller 0 5 10  x/D  15  20  430 433 437 432  Mg (%)  Mg (%)  (a)  80 EB/S 60  S  40  B  B  EB  B  EB  EB EB  Usl(m/s) N (rpm)  B/EB  20 Impeller 10  25  0  5  10  x/D  EB  15  0.5 1.0 2.0 3.0  430 425 440 441  20  25  (c) 100  Mg (%)  80 EB/S 60  S  EB EB EB  B  B  40  EB EB EB  B  B/EB  Usl(m/s) N (rpm) 0.5 1.0 2.0 3.0  20 Impeller 0  5  10  x/D  15  427 432 447 442  20  25  Figure D.20: Gas mixing index and flow regimes as a function of dimensionless  distance downstream for Cm = 3.0%, virtually identical impeller speeds and various superficial liquid/pulp velocities at: (a) Usg = 0.22 m/s; (b) Usg = 0.33 m/s; (c) Usg = 0.44 m/s. Letters identify flow regimes as in Figure 6.6.  299  

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