Residence time packed catalytic reactor

More complex reactors, like packed-bed reactors or catalytic monoliths, consist of many physically separated scales, with complex nonlinear interactions between the processes occurring at these scales. Figure 3 illustrates scale separation in a packed-bed reactor. The four length (and time) scales present in the system are the reactor, catalyst particle, pore scale, and molecular scale. The typical orders of magnitude of these four length scales are as follows reactor, lm catalyst particle, 10 2m (1cm) macropore scale, 1pm (10 6m) micro-pore/molecular scale, 10 A (10 9m). The corresponding time scales also vary widely. While the residence time in the reactor varies between 1 and 1000 s, the intraparticle diffusion time is of the order of 0.1s and is 10 5s inside the pores. The time scale associated with molecular phenomenon like adsorption is typically less than a microsecond and could be as small as a nanosecond. 

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. 

This is the first reactor reported where the aim was to form micro-channel-like conduits not by employing microfabrication, but rather using the void space of structured packing from smart, precise-sized conventional materials such as filaments (Figure 3.25). In this way, a structured catalytic packing was made from filaments of 3-10 pm size [8]. The inner diameter of the void space between such filaments lies in the range of typical micro channels, so ensuring laminar flow, a narrow residence time distribution and efficient mass transfer. 

CSTR for most reactions. These conditions are best met for short residence times where velocity profiles in the tubes can be maintained in the turbulent flow regime. In an empty tube this requires high flow rates for packed columns the flow rates need not be as high. Noncatalytic reactions performed in PFRs include high-pressure polymerization of ethylene and naphtha conversion to ethylene. A gas-liquid noncatalytic PFR is used for adipinic nitrile production. A gas-solid PFR is a packed-bed reactor (Section IV). An example of a noncatalytic gas-solid PFR is the convertor for steel production. Catalytic PFRs are used for sulfur dioxide combustion and ammonia synthesis. 

Catalytic oxidation reactions were carried out in a conventional fwed bed reactor under atmospheric pressure [10]. The flow rate through the reactor was set at 500 cm min and the gas hourly space velocity (GHSV) was set at 15000 h. The residence time based on the packing volume of the catalyst was 0.24 s. Following the reactor, a portion of the effluent stream was delivered and analysed on-line using a Hewlett Packard 5890 Series II gas chromatograph (GC) equipped with an electron capture detector (ECD) and a thermal conductivity detector (TCD), and controlled with HP ChemStation software. The concentration of the chlorinated feeds was determined by the ECD after being separated in a HP-VOC column. 

In conventional fixed-bed reactors, catalyst particles of various sizes are often randomly distributed, which may lead to inhomogeneous flow patterns. Near the reactor walls, the packing density is lower than the mean value, and faster flow of the fluid near the wall is unavoidable. As a result, reactants may bypass the catalyst particles, and the residence time distribution (RTD) will be broadened. Moreover, the nonuniform access of reactants to the catalytic surface diminishes the overall reactor performance and can lead to unexpected hot spots and even to reactor runaway in the case of exothermic reactions. 

In Fig. 5.18 IMRCF (s=l) corresponds to the PBMR with catalyst pellets with a non-uniform catalyst distribution, while CMR ( =1) is the CMR with the catalyst placed non-uniformly on the membrane surface in contact with the catalytic reactor feed. IMRCF (a(s)=I) and CMR (a( )=I) correspond to the PBMR and CMR with uniform catalyst distributions. The conventional packed-bed reactor (FBR in Figure 5.18) exhibits conversions, which are below the equilibrium conversion, and for large residence times are lower than those exhibited by the CMR and the PBMR. The highest conversions are obtained with the non-uniform activity (Dirac delta case) profiles. This result was explained on the basis that the access of the reactants to the active catalytic sites was not limited by diffusion. When the catalyst is uniformly distributed the PBMR exhibits better performances than the CMR. It is interesting to note that at low residence times the packed-bed reactor conversion is higher than that of the PBMR with a uniformly distributed catalyst this is because in this case for the PBMR the reactants are only partially in contact with the catalyst due to diffusional limitations. 

Before the use, the catalyst has been pressed, crushed and sieved to obtain a 400-700 pm fraction. N 0 inlet concentration to the catalytic reactor has been 300 or 600 ppm, the carrier gas being He. The effect of the presence of O in the feed has been also investigated in the range from 100 to 5000 ppm. The catalytic bed has been packed with 1.43 g of Cu-ZSM5 catalyst, whereas a 37.5 Nl/h total flow rate for the reacting gases has been used, corresponding to a residence time of 1.4-10 min... 

If there is only one chemical reaction on the internal catalytic surface, then vai = — 1 and subscript j is not required for all quantities that are specific to the yth chemical reaction. When the mass transfer Peclet number which accounts for interpellet axial dispersion in packed beds is large, residence-time distribution effects are insignificant and axial diffusion can be neglected in the plug-flow mass balance given by equation (22-11). Under these conditions, reactor performance can be predicted from a simplified one-dimensional model. The differential design equation is... 

Calculate the average residence time for the packed catalytic tubular reactor in terms of the effectiveness factor, and ultimately the intrapellet Damkohler number, without complications due to residence-time distribution effects. 

Problem. Think about the overall strategy that must be implemented to account for the effect of interpellet axial dispersion on ihe outlet concentration of reactant A when Langmuir-Hinshelwood kinetics and Hougen-Watson models are operative in a packed catalytic tubular reactor. Residence-time distribution effects are important at small mass transfer Peclet numbers. 

This dimensionless parameter is required to analyze residence-time distribution effects in packed catalytic tubular reactors, and it corresponds to the quantity on the horizontal axes of Figures 22-2 and 22-3, which compare ideal vs. non-ideal reactor performance. 

Gas chromatography is a separation technique based on the fact that different components in the mixture exhibit different average residence times due to interactions with the porons packing material. These interactions can be classified as intrapellet diffusion and the column operates similar to a packed catalytic tubular reactor. The important mass transfer mechanisms are convection and diffusion. Hence, it is important to calculate the mass transfer Peclet number that represents an order-of-magnitude ratio of these two mass transfer rate processes. Intrapellet diffusion governs residence times, and interpellet axial dispersion affects the degree to which the output curve is broadened. For axial dispersion in packed columns and packed catalytic tubular reactors. 

Step 21. Calculate the time constant for convective mass transfer through the packed catalytic tubular reactor in units of minutes, which is equivalent to the residence time ... 

At high-mass-transfer Peclet numbers, sketch the relation between average residence time divided by the chemical reaction time constant (i.e., r/co) for a packed catalytic tubular reactor versus the intrapeUet Damkohler number Aa, intrapeiiet for zeroth-, first-, and second-order irreversible chemical kinetics within spherical catalytic pellets. The characteristic length L in the definition of Aa, intrapeiiet is the sphere radius R. The overall objective is to achieve the same conversion in the exit stream for all three kinetic rate laws. Put all three curves on the same set of axes and identify quantitative values for the intrapeiiet Damkohler number on the horizontal axis. 

Calculate the conversion of reactant A in the exit stream of this packed catalytic tubular reactor for the residence times t given below ... 

The drawback of randomly packed microreactors is the high pressure drop. In multitubular micro fixed beds, each channel must be packed identically or supplementary flow resistances must be introduced to avoid flow maldistribution between the channels, which leads to a broad residence time distribution in the reactor system. Initial developments led to structured catalytic micro-beds based on fibrous materials [8-10]. This concept is based on a structured catalytic bed arranged with parallel filaments giving identical flow characteristics to multichannel microreactors. The channels formed by filaments have an equivalent hydraulic diameter in the range of a few microns ensuring laminar flow and short diffusion times in the radial direction [10]. 

The catalytic filaments were introduced into the tubular reactor in the form of threads. A bundle of 100 filaments with a diameter of 7 pm each formedthreads of diameter of about 0.5 mm. The catalytic threads were placed in parallel into the tube to form a cylindrical catalytic bed of several centimeters length. This arrangement gives about 300 threads per cm within the tube cross section with a porosity of = 0.8. The specific surface per volume is in the order of 10 m m and, thus, about 50 times higher compared to washcoated tubes of the same inner diameter [8]. The performance comparison under identical experimental conditions with randomly packed beds with particles of silica and y-alumina of different shapes and sizes showed significantly broader residence time distribution compared to the structured filamentous packing with about five times lower pressure drop for the same hydraulic diameter and comparable gas flow rates. 

To combine the advantages of packed-bed and catalytic wall microreactors, catalytic bed microreactors were proposed recently. In this novel reactor design, the catalyst is applied on metallic filaments or wires which are incorporated in a microreactor, leading to a low pressure drop and a nanow residence time distribution [87-89]. By insertion of metallic wires a uniform gas distribution and a reduced risk of temperature gradients is obtained. However, similarly to catalytic wall microreactors, an increase in the specific surface area of the grid or wire is required. In addition to metallic wires and grids, modified ceramic tapes can also be used as a catalyst support [90]. 

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