Fluid catalysts measurement

In LS systems the mixing is less problematic and the catalyst is not necessarily held in a packed bed, but can be suspended in the fluid with measures to prevent washing out of the reactor and accumulation at the exit. 

Porous ceramic or micrometallic filters are very effective for recovering entrained fines from gas streams (228,252). Multiple installations are required because it is necessary to blow back each filter element periodically to dislodge the catalyst cake that builds up on the surface and leads to increased pressure drop. Filters have been used for catalyst recovery in other fluid-catalyst processes where high cost or other considerations justify extraordinary measures to minimize catalyst losses. However, this expedient has not been employed in commercial fluid cracking units because losses are readily controlled to a reasonable level by simpler means. In fact, intentional discard of catalyst is often practiced, in addition to normal losses, in order to maintain catalyst quality at a high level by permitting increased additions of fresh catalyst. 

Currently available data for the flow properties of the fluidized catalyst bed are fragmentary, since the local motion of the emulsion phase is diflicult to measure experimentally. Therefore, it is useful to clarify the flow properties of the bed in terms of our knowledge of bubble columns. First, the fluid-dynamic properties of the bubble columns will be explained then, the available data will be adapted to apply to fluid catalyst beds. The reader will be able to picture an emulsion phase of carefully prepared catalyst particles operating in intense turbulence for fluidized beds under conditions of practical interest. This turbulence distinguishes the flow properties of fluid catalyst beds from those of widely studied teeter beds. 

Inside the catalyst pores, the composition of the gas is at the equilibrium composition shown above. The values of PcmlPco and Pm IPcm of the bulk fluid and the equilibrium composition are greater than the values fliased on Ni catalyst) measured for (Boudouard reaction) and Kg (methane cracking reaction) ... 

Feed to the reactor is a simulated crude methanol stream consisting of a mixture of 83% methanol and 17% water. The reactor accommodates a dense fluid catalyst bed measuring 2 ft in diameter by 40 ft in hei t. The feed can either be injected as a liquid or vaporized and superheated before entering the reactor. The feed passes throu the bed, where it is converted quantitatively... 

The measurement of the linear velocity as a function of shaft RPM can be done at room temperature and pressure in air. It is best to do this on the catalyst already charged for the test. Since u is proportional to the square of the head generated, the relationship will hold for any fluid at any MW, T, and P if the u is expressed at the operating conditions. The measurement can be done with the flow measuring attachment and flow meter as shown in Figure 3.5.1. 

Temperature gradient normal to flow. In exothermic reactions, the heat generation rate is q=(-AHr)r. This must be removed to maintain steady-state. For endothermic reactions this much heat must be added. Here the equations deal with exothermic reactions as examples. A criterion can be derived for the temperature difference needed for heat transfer from the catalyst particles to the reacting, flowing fluid. For this, inside heat balance can be measured (Berty 1974) directly, with Pt resistance thermometers. Since this is expensive and complicated, here again the heat generation rate is calculated from the rate of reaction that is derived from the outside material balance, and multiplied by the heat of reaction. 

Here Iq is the thermal conductivity of the system, consisting of the porous solid and the reacting fluid inside the pores. This is the most uncertain value, while everything else is measurable. Two things must be remembered. First, data on thermal conductivity of catalysts are approximate. The solid fraction of the catalyst (1-0) always reduces the possibility for diflhision, while the solid can contribute to the thermal conductivity. Second, the outside temperature difference normal to the surface or Daiv, will become too high, much before the inside gradient can cause a problem. See Hutching and Carberry (19), Carberry (20). 

Measurements of the true reaction times are sometimes difficult to determine due to the two-phase nature of the fluid reactants in contact with the solid phase. Adsorption of reactants on the catalyst surface can result in catalyst-reactant contact times that are different from the fluid dynamic residence times. Additionally, different velocities between the vapor, liquid, and solid phases must be considered when measuring reaction times. Various laboratory reactors and their limitations for industrial use are reviewed below. 

Ross (R2) measured liquid-phase holdup and residence-time distribution by a tracer-pulse technique. Experiments were carried out for cocurrent flow in model columns of 2- and 4-in. diameter with air and water as fluid media, as well as in pilot-scale and industrial-scale reactors of 2-in. and 6.5-ft diameters used for the catalytic hydrogenation of petroleum fractions. The columns were packed with commercial cylindrical catalyst pellets of -in. diameter and length. The liquid holdup was from 40 to 50% of total bed volume for nominal liquid velocities from 8 to 200 ft/hr in the model reactors, from 26 to 32% of volume for nominal liquid velocities from 6 to 10.5 ft/hr in the pilot unit, and from 20 to 27 % for nominal liquid velocities from 27.9 to 68.6 ft/hr in the industrial unit. In that work, a few sets of results of residence-time distribution experiments are reported in graphical form, as tracer-response curves. 

Strictly gas-phase CSTRs are rare. Two-phase, gas-liquid CSTRs are common and are treated in Chapter 11. Two-phase, gas-solid CSTRs are fairly common. When the solid is a catalyst, the use of pseudohomogeneous kinetics allows these two-phase systems to be treated as though only the fluid phase were present. All concentration measurements are made in the gas phase, and the rate expression is fitted to the gas-phase concentrations. This section outlines the method for fitting pseudo-homogeneous kinetics using measurements made in a CSTR. A more general treatment is given in Chapter 10. 

While the previously described techniques were measuring the nanoscopic and microscopic properties of the catalyst pellets, respectively, fluid transport within... 

In addition to flow, thermal, and bed arrangements, an important design consideration is the amount of catalyst required (W), and its possible distribution over two or more stages. This is a measure of the size of the reactor. The depth (L) and diameter (D) of each stage must also be determined. In addition to the usual tools provided by kinetics, and material and energy balances, we must take into account matters peculiar to individual particles, collections of particles, and fluid-particle interactions, as well as any matters peculiar to the nature of the reaction, such as reversibility. Process design aspects of catalytic reactors are described by Lywood (1996). 

In these studies, chemical conversion was determined in situ by measuring the lH resonance associated with OH groups present. In practice two such resonances exist associated with chemical species inside and outside the catalyst particles, respectively. The difference in chemical shift between these intra- and inter-particle species arises because of the different electronic environment of the molecules inside the catalyst particles compared to their environment in the bulk fluid in the inter-particle space. In this work, chemical conversion was determined from the MR signal acquired from species in the inter-particle space of the bed because the signal from inside the catalyst particles is also going to be influenced, to an unknown extent, by relaxation time contrast. In addition to possible relaxation contrast effects, there will also be modifications to the chemical shifts of individual species resulting from adsorption onto the catalyst this may cause peak broadening and reduces the accuracy with which we can determine the chemical shift of the species of interest. As follows from eqn (11) which describes the esterification reaction of methanol and acetic acid to form methyl acetate and water ... 

Whilst it is obviously valuable to measure the solubility of reagents in the SCF, it is important to be aware that the solubility in a multicomponent system can be very different from that in the fluid alone. It is also important to note that the addition of reagents and catalysts can have a profound effect on the critical parameters of the mixture. Indeed, at high concentrations of reactants, the mole fraction of C02 is necessarily lower and it may not be possible to achieve a supercritical phase at the temperature of interest. Increases in pressure (i.e. further additions of C02) could yield a single liquid phase (which would have a much lower compressibility than scC02). For example, the Diels-Alder reaction (see Chapter 7) between 2-methyl-1,3-butadiene and maleic anydride has been carried out a pressure of 74.5 bar and a temperature of 50 °C, assuming that this would be under supercritical conditions as it would if it were pure C02. However, the critical parameters calculated for this system are a pressure of 77.4bar and a temperature of 123.2 °C, far in excess of those used [41]. 

The pellets leave a fraction e unoccupied as they pack into the reactor so the fraction 1 — is occupied by the catalyst. The pellet is usually porous, and there is fluid (void space) both between catalyst pellets and within pellets. We measure the rate per unit area of pellet of assumed geometrical volume of pellet so we count only the void fraction external to the pellet. 

Activity measurements. Activity and selectivity measurements were performed at 10 psig in a 14-mm internal diameter glass fluid bed reactor using 25 grams of 90 to 38 micron catalyst particles. A reactant mixture of approximately 18 volume % 02 7 volume X NH3 and 7 volume % CH3OH and the balance of helium was fed to the catalyst, and temperature and contact time were varied to find the optimum yield of HCN. Optimum reaction temperatures were found to range from 425° to 475°C with contact times of 3 to 5 seconds (calculated at STP>. Fixed bed reactor studies produced similar results. The yields reported in this paper are based on carbon fed, unless otherwise noted. More details on catalyst performance can be found in our patents (7,8). 

Hydrothermal (steam) stability is also important, in as much as the catalyst must pass through a high temperature stripping zone in which the usual fluid stripping medium is steam. In our laboratory, zeolite hydrothermal stability is measured by comparing the x-ray crystallinity of the unknown faujasite sample with that of a fully rare earth exchanged reference standard following a 3 hour, 100% steam, 1500 F treatment. 

The example furthermore shows that diffusion from the bulk fluid phase toward the volume near the IRE, which is probed by the evanescent field, has to be accounted for because it may be the limiting step when fast processes are investigated. The importance of diffusion is more pronounced when a catalyst layer is present on the IRE, because of the diffusion in the porous film is much slower than that in the stagnant liquid film. Indeed, the ATR method, because of the measurement geometry, is ideally suited to characterization of diffusion within films (50,66-68). Figure 16 shows the time dependence of absorption signals associated with cyclohexene (top) and i-butyl hydroperoxide (TBHP, bottom). Solutions (with concentrations of 3mmol/L) of the two molecules in cyclohexane and neat cyclohexane were alternately admitted once to... 

---