Graduation date: 1994

Thesis (Ph. D.)--Technische Universiteit Delft, 1997. / Summary and acknowledgements in Dutch.

Thesis (Ph. D.)--Technische Universiteit Delft, 1997. / Summary and acknowledgements in Dutch.

Thesis (doctoral)--Technical University of Delft, 2007. / "Gas-cooled fluidized bed nuclear reactor ... (FLUBER)"--P. 3. Includes bibliographical references (p. [139]-148).

The effect of particle size distribution (PSD) on the performance of a fluidized bed reactor was investigated using the ozone decomposition reaction, combined with the study of hydrodynamics, for fresh and spent fluid cracking catalysts, each having three particle size distributions - wide, narrow and bimodal - all with nearly the same mean diameter (60 µm), the same particle density and the same BET surface area. The superficial gas velocity was varied from 0.1 to 1.8 m/s to include the bubbling, slugging, turbulent and fast fluidization regimes. The catalytic rate constant, based on the volume of the particles, ranged from 2 to10 s⁻¹, while the static bed height was varied from 0.15 m to 1 m. Four different multi-orifice gas distributors with different hole diameters (2.2 to 5.1 mm) and hole numbers (4 and 21) were also tested to evaluate the influence of gas distributor on the performance of fluidized bed reactors. The particle size distribution was found to play a larger role at higher gas velocities than at lower velocities. At low gas velocities (Uf ≤ 0.2 m/s), the reaction conversion was not greatly affected by the PSD. However, with an increase in gas velocity the PSD effect became larger. The wide size distribution gave the highest reactor efficiency, defined as the ratio of the volume of catalyst required in a plug flow reactor to that required in the fluidized bed reactor to achieve the same conversion, while the narrow blend gave the lowest. The differences are not solely a function of the "fines content". The influence of particle size distribution on the hydrodynamics of fluidization was evaluated by measuring particle concentrations in voids, bubble sizes, and dense phase expansion. When the superficial gas velocity exceeded 0.1 m/s, the bed with the wide size distribution usually gave the highest particle concentration inside the voids, the smallest bubble size and the greatest dense phase expansion at the same operating conditions. There is evidence that there is a greater proportion of "fines" present in the voids than in the overall particle size distribution. This has been explained in terms of the throughflow velocity inside bubbles being of the same order as the terminal velocity of typical "fines", causing these particles to spend longer periods of time inside the voids. The effect of the PSD on the fluidization regime and its transitions was determined by measuring pressure fluctuations along the column. The earliest transition from bubbling or slugging to turbulent fluidization occurred in the bed of wide size distribution, while the latest corresponded to the narrow PSD. For particles of wide size distribution, higher conversion was achieved for the turbulent and fast fluidization regimes than for the bubbling fluidization regime under otherwise identical conditions, while for particles of narrow size distribution, the dependence of conversion on regime was small. Hence, for reactors of wide PSD, the performance can be improved significantly by operating in the turbulent or fast fluidization regime, while for particles of narrow size distribution, the benefit of operating at high gas velocity is slight at best. The PSD influence should be considered in modelling fluidized bed reactors. The "Two-Phase Bubbling Bed Model" has been modified to account for PSD effects. For the reactor of wide particle size distribution operated at high gas velocities, a single-phase axial dispersion model with closed inlet and open outlet boundary conditions appears to be suitable to predict the performance. It was also found that a high pressure drop across the gas distributor was not sufficient to maintain good performance of the distributor. The reactor efficiency in the entry region was higher for a distributor with a greater number of orifices, even though it had a lower pressure drop, than for a distributor plate with fewer larger holes. / Applied Science, Faculty of / Chemical and Biological Engineering, Department of / Graduate

No description available.

Call number: LD2668 .T4 1984 N46 / Master of Science

Sasol's Research and Development Division has identified several proprietary gas-liquid reactions where very low superficial gas velocities « 0.8 cmls) are required to obtain desired conversions in a bubble column reactor. A review of existing literature has shown that research in bubble column reactors is typically conducted in the superficial gas velocity range of 1 - 40 cmls. Traditionally bubble column reactors are designed via the application of empirical correlations which are only valid under specific conditions. There is a danger of under or over design if incorrect nonadjustable parameters such as liquid dispersion coefficients, mass transfer coefficients and gas hold-up values are used. To this extent, a hydrodynamic study was undertaken at superficial gas velocities lower than 0.8 cmls, to determine whether existing correlations are valid in this little investigated superficial gas velocity regime. Three bubble column reactors were designed and set up to perform hydrodynamic studies: • 22 cm inner diameter QVF glass column, 190 cm tall • 30 cm inner diameter 304 stainless-steel column, 200 cm tall • 30 cm inner diameter QVF glass column, 80 cm tall All measurements were undertaken in an air/water system. Gas hold-up measurements revealed that at the investigated gas flow rates, the gas hold-up was less than 1 % and as such was not investigated extensively. Partition plates were installed into the bubble columns and residence time distribution measurements were undertaken. The bubble columns were found to behave identically to the well known tanks in series model (Levenspiel, 1962). Liquid dispersion coefficients were measured via two methods. Batch liquid measurements were undertaken via the method of Ohki and Inoue (1970) and continuous liquid residence time distribution measurements were also undertaken. Data reduction was performed for both methods using the axial dispersion model to regress the liquid dispersion coefficient EL_ Both methods yielded equivalent results. The effect of distributor plate geometry on EL was also investigated and proved not to affect EL. It was found that existing literature correlations developed at higher superficial gas velocities failed to accurately predict the measured dispersion coefficients obtained in this study_ Correlation of the EL values with column diameter and superficial gas velocity showed EL to be a weak function of diameter as compared to existing correlations. This will have a significant effect on scale-up to larger column diameters. / Thesis (M.Sc.Eng.)-University of KwaZulu-Natal, Durban, 2005.

Bibliography: leaves 210-223.

This work is concerned with the effect of vessel geometry on the hydrodynamics of fluidisation of a bed of milled iron oxide. The effect of going from a cold model representative of a typical pilot plant reactor to one simulating a semi-commercial unit is quantified, and various reactor internal configurations on the latter are evaluated. The experimental approach is one based on residence time testing and model fitting with parameter optimisation. A model screening aimed at identifying the most reasonable modelling approach is included, and altogether seven models in two categories are formulated and solved in the dynamic mode. Three of these models are considered novel at present, along with the dynamic solutions to two of the others. The residence time technique involves methane as an inert tracer in air, and continuous analysis of gas withdrawn from the bed via sample probes by a pair of flame ionisation detectors. The process stimulus is governed by a pseudo-random binary sequence, and correlation analysis is employed for noise reduction. A Fourier transform routine, developed from first principles, converts a pair of correlation functions to a process frequency response, and model predictions are compared with the experimental data in this form. Two parameters per model are fitted, and the residual error at the optimum parameter combination provides a means of identifying the best-fitting model. The optimised parameters of this model are regarded as estimates of those of the actual process. Five models compete in the first screening category. Four of these have appeared in the literature in one form or another, and the fifth is novel in that it accounts for axial mixing in the bubble phase by employing multiple plug flow units. This model, referred to as the multiple bubble-track or MBT model, is shown to fit the experimental data better than any of the other models in both bubbling and slugging systems. This suggests that employing multiple plug flow units in parallel for the bubble phase is mechanistically more correct than employing a single plug flow unit. The second screening category is related to the situation in which gas is sparged into an already fluidised bed at some height above the main distributor. The two models in this category are both considered novel, and describe opposite extremes of possible behaviour in one particular sense: one assumes rapid coalescence between grid and sparger bubbles, and the other none at all. The laterally segregated bubble phase or LSBP model emerges as the better process description.The formulation of this model suggests that physically, bubbles from the sparger tend to retain their identity as they pass through the bed. Crossflow ratios estimated on the basis of the best-fitting model in each category point to the existence of a very strong scale-up effect. From the shape of the crossflow profiles it appears that most of the interphase mass transfer occurs in the bottom meter or so of the bed, and it is suggested that grid design is the most significant controlling factor. The presence or otherwise of vertical coils in the bed is shown to have no significant effect on crossflow, and mass transfer between sparger bubbles and the dense phase is shown to be similar to that between grid bubbles and the dense phase. Finally, it is demonstrated that the axial crossflow profile in the bubbling bed is consistent with the concept of an axially invariant mass transfer coefficient based on bubble to dense phase interfacial area. / Thesis (Ph.D.)-University of Natal, Durban, 1984.