Regular and irregular catalyst packings are extensively used in the chemical industry for promoting mass transfer and chemical reaction between gas and liquid. Traditionally, randomly packed beds of catalyst particles operating in co-current downflow operation, often termed trickle-bed reactors, have been used because of the high mass-transfer rates achievable. More recently, structured supports are increasingly considered for use because of the potential improvements they offer with respect to, for example, the decoupling of heat-and mass-transfer phenomena, operation under reduced pressure drop conditions and at much higher gasrliquid Ž flow rates, and a greater resistance to attrition Irandoust and Andersson, 1988a,b; Irandoust et al., 1998; Kapteijn et al., . 1999 . Monoliths, which comprise a metal or ceramic structure with a large number of straight or parallel channels, are an example of such structured supports, and their use in solid-catalyzed gas-phase chemical reactions is well established. One example is the monolithic exhaust converter used throughout the automotive industry. In contrast, the application of monoliths to gas-liquid reactions is not well advanced and a significant research activity exists to provide the necessary understanding to enable reliable scale-up for process op-Ž . eration Nijhuis et al., 2001 . Undoubtedly, such efforts would be significantly aided if it were possible to develop an in-situ probe of the multiphase transport and reaction phenomena occurring within these porous structures. Magnetic resonance techniques show particular promise in this regard because of their ability to provide chemically-specific information on both the internal phase distribution and transport processes Ž . occurring within three-dimensional 3-D optically opaque systems.Visualization of gas-liquid flow in ceramic monoliths has Ž . previously been attempted by Mewes et al. 1999 using the capacitance tomography technique. While temporal resoluCorrespondence concerning this article should be addressed to L. F. Gladden. Ž tion of capacitance tomography is high 100 images per sec-. ond can be achieved , spatial resolution is relatively low with typical in-plane resolutions being 5᎐10% of the diameter of the system under investigation, thereby making impossible visualization of phase distribution within a single channel. In Ž . contrast, magnetic resonance imaging MRI has the poten-Ž . tial for relatively high spatial resolution say, 30᎐200 m , but the temporal resolution obtained for gas-liquid flow has to date been far too slow to visualize dynamic processes occurring in reactors. Time-averaged visualizations of singleand two-phase flow have been reported with data acquisition Ž times of minutes to hours Sederman et al., 1998;Tallarek et al. 1998;Johns et al., 2000; Sederman and Gladden, 2001; . Mantle et al., 2001 . Fast velocity imaging techniques have been reported which are able to acquire liquid velocity im-Ž ages in several minutes Seifert et al., 2000; Scheenen et al., . 2000 , but ...

No abstract

The hydrogenation of cinnamaldehyde was investigated using a 5% Pd/C catalyst in a 250 cm3 stirred tank reactor and 500 cm3 autoclave. The experiments were carried out at 273–343 K and 0·1–1·1 MPa. Non‐polar solvents, e.g. toluene, decane, methylcyclohexane, decalin, ether and heptane, and polar solvents such as methanol, ethanol, propan‐1‐o1, propan‐2‐ol, butan‐1‐ol and butan‐2‐ol were used to study the selectivity with respect to hydrocinnamaldehyde formation, the reaction kinetics and mass transfer. The additives, such as potassium acetate, ferrous chloride, ferrous sulphate and quinoline were incorporated into the catalyst in order to improve the catalyst selectivity, which was observed especially in the case of potassium acetate. © 1998 SCI

No abstract

The Cocurrent Downflow Contactor (CDC) has been studied as a three phase reactor and in addition, chemically enhanced mass transfer studies have confirmed the very high gas hold-up values previously indicated by photographic methods ( E~ = 0.5-0.6). Mass transfer measurements for the O,/H,O system have shown that volumetric mass transfer coeficients (kLa) are in the range 0.25-0.3 s -' for unpacked and packed CDC reactors, while Coz+ ion-catalysed sulphite oxidation exhibited enhanced kLa values of 0.55 s-' with interfacial areas in the range of 1000-6000 mz m-3. A model first order reaction, the palladium-catlysed hydrogenation of itaconic acid, was examined both in a small stirred batch reactor and in the CDC. Low degrees of mass transfer resistances were observed (both gas-liquid and liquid-solid) especially in the case of the CDC when used as a slurry and fixed bed reactor, with liquid-solid mass transfer resistances being the the range 1-10%. This was confirmed by energy of activation measurements in the range 30-45 kJ mol-'. The CDC was used in slurry and fixed bed form for the respective hydrogenation of the triglycerides, rapeseed and soyabean oils. The reaction was predominantly surface reaction rate controlled with energies of activation in the range 47-58 kJ mol-', using palladium and nickel catlysts. Reaction selectivities were high, especially in respect of linolenate removal and the fixed bed CDC was slightly superior to the slurry reactor.Key words: catlytic hydrogenation, three-phase reactors. * To whom correspondence should be addressed. NOTATION PH: 55Concentration of hydrogen at the catalyst surface (mol m-3) Equilibrium concentration of gas at the gasliquid interface (mol m-3) Diameter of catalyst particle (m) Diffusivity of hydrogen (m2 s -') Liquid energy dissipation (g cm-2 s -') Energy of activation (kcal mol-or kJ mol-') Liquid flow rate in column (m3 s-') Height of dispersion (m) Volumetric gas-liquid mass transfer coefficient Pressure drop per unit bed height in a fixed bed reactor (dyne cm-3) Bed voidage (dimensionless) Liquid voidage (dimensionless) Catalyst effectiveness factor (dimensionless) Viscosity of the liquid (kg ms-') Density of the liquid (kg m-3) Density of the catalyst particle (kg m-3)

m e oxidation of aqueous solutions of phenol as a typical model pollutant has been carried out in the presence of an ultra-violet (W) irradiated TiO, catalyst in a CDC reactor. The CDCR was fitted with internally mounted 30 W and 1.0 kW W lamps. The reactions were carried out at 40-50 "C and 202.6 P a , with the reactor being operated in closed loop recycle mode and suspended 2atabst being recirculated. The CDCR is a device of high mass transfer efliciency, giving unusually large gas hold-up (approximately 50%). The CDCR was operuted (19 with oxygen mass transfer and dissolution in the zone above the Wsource and (id with oxp gen dispersion and mass transfer occurring along and around the lamp housing. Using the higher powered lamp, lOP? conversion of phenol was obtained from solutions containing 100 mg/dm3. Under spec@ conditions the presence of TiO, catalysts was obserued to give the most rapid oxidation degradation of thephe nol.

Aspergillus fumigatus was cultured in disc-turbine-agitated vessels and in an air-lift fermentor. In the agitated vessels the yield of cellulase was reduced when the agitation rate was increased, although extracellular protein levels rose. The enzyme complex itself was shown to be exceptionally stable under conditions similar to those in the agitated vessels, so probably shear damage to the mycelium had occurred, liberating intracellular contents. These appeared to contain an inhibitor that could be removed by fabricated inorganic protein absorbents, such as kieselguhr and alumina. However, the inhibitor was not likely to be protease, since only relatively low levels could be detected and its identity has not been established. The use of an air-lift fermentor avoided the shear effects due to use of the disc turbine agitator in the conventional fermentors, and yields of enzyme were then found to increase by about 20%, maximum yields being obtained at maximum K(L)a values.

The hydrogenation of the triglyceride oil, soya bean oil, has been studied in the temperature range 130±160°C and in the pressure range 100±600 kPa using (i) a 5% w/w Pd/C slurry catalyst and (ii) a 3% w/w Pd/Al 2 O 3 Raschig ring catalyst in a cocurrent down¯ow contactor (CDC) reactor. Separate studies of residence time distribution (RTD) were carried out in a modi®ed CDC device in order to determine dispersion numbers and dispersion coef®cients. The RTD measurements indicated that the overall¯ow was a mixture of well-mixed and plug¯ow for the unpacked CDC, so that the entry section (0±30 cm from entrance) was perfectly mixed and the remainder of the column (30±130 cm) gave predominantly plug¯ow behaviour. The introduction of random packing in the form of 13 mm Raschig rings gave rise to increased back mixing in the lower part of the CDC and the overall dispersion number increased due to liquid and gas circulation around the packing elements. Kinetic studies revealed an initial rate reaction order of 1.24±1.26 with respect to hydrogen concentration both in slurry and ®xed bed CDC reactors and is interpreted as a combination of a parallel pair of ®rst and second order reactions during the initial stages of reaction. Mass transfer coef®cients for gas absorption (k L a) and liquid±solid mass transport (k s ) were determined for both types of reactor. The k L a values lay in the range 1.0±3.33 s À1 and the liquid±solid transport resistances (X LS ) were all`1%, so that the reaction was almost totally surface reaction rate controlled. Apparent energy of activation measurements gave values of E A =49 AE6 kJ mol À1 , which is strongly indicative of surface reaction rate control involving the hydrogenation of an ole®nic double bond. The selectivity in respect of linolenate (three double bonds) removal and linoleate (two double bonds) retention was high with, for palladium, relatively low trans-isomer production (`30%). The overall selectivity was slightly, but signi®cantly, better for the ®xed bed CDC reactor and this is attributed to the greater degree of plug¯ow behaviour in the latter, despite the bed causing an increase in dispersion number. However, there is no reaction in the well-mixed section of the ®xed bed CDC reactor as there is in the slurry CDC reactor and this is likely to improve selectivity in a consecutive reaction sequence. NOTATION aGas±liquid interfacial area (m À1 ) a p External particle area (m À1 ) A Column cross-sectional area (m 2 ) C Aig Gas-side concentration of A at gas±liquid interface (mol m À3 ) C Ae Equilibrium concentration of A in solution (mol m À3 ) C As Surface concentration of A (mol m À3 ) d c Column diameter (m) d p Particle diameter (m) D Dispersion coef®cient (cm 2 s À1 ) D A Diffusion coef®cients of A (H 2 ) in solvent (m 2 s À1 ) E l Liquid energy dissipation term (kgf m À2 s À ) H dispPFR Height of PFR section (m) H p Height of packed bed (m) H s Height of single phase section (m) IV Iodine value k L Gas±liquid mass transfer coef®cients (m s À1 ) k r Reaction rate constant (m...