Microreactor tubular reactors

Quite new ideas for the reactor design of aqueous multiphase fluid/fluid reactions have been reported by researchers from Oxeno. In packed tubular reactors and under unconventional reaction conditions they observed very high space-time yields which increased the rate compared with conventional operation by a factor of 10 due to a combination of mass transfer area and kinetics [29]. Thus the old question of aqueous-biphase hydroformylation "Where does the reaction takes place " - i.e., at the interphase or the bulk of the liquid phase [23,56h] - is again questionable, at least under the conditions (packed tubular reactors, other hydrodynamic conditions, in mini plants, and in the unusual,and costly presence of ethylene glycol) and not in harsh industrial operation. The considerable reduction of the laminar boundary layer in highly loaded packed tubular reactors increases the mass transfer coefficients, thus Oxeno claim the successful hydroformylation of 1-octene [25a,26,29c,49a,49e,58d,58f], The search for a new reactor design may also include operation in microreactors [59]. 

The microreactor technique described by the authors offered 30 and 50 times higher STY values than the conventional tubular reactor and the slurry reactor, respectively, suggesting microreactors to be a successful technique in mass transfer controlled reactions. 

Now we shall discuss the method used to calculate the "cup"-averaged MWD-H, in which all portions of a polymerized liquid are mixed and averaged in a "cup" (vessel) positioned after the reactor. In this analysis, recourse was made to the so-called "suspension" model of a tubular reactor. According to this model, the reaction mass is regarded as an assemblage of immiscible microvolume batch reactors. Each of these microreactors moves along its own flow line. The most important point is that the duration of the reaction is different in each microreactor, as the residence time of each microvolume depends on its position at any given time, i.e., on its distance from the reactor axis. 

Methane decomposition experiments were conducted in a 5.0 ml fixed bed quartz microreactor using 0.3 g of catalysts. The catalysts were arranged within the reaction zone in several layers separated with ceramic wool to prevent clogging of the reactor due to produced carbon. The reactor temperature was maintained constant via a type K thermocouple and Love Controls microprocessor. The tubular reactor was made out of alumina and quartz tubings (I.D. 3-6 mm). 

The microreactor used for evaluating catalyst activity was operated at atmospheric pressure and consisted of Barter manual mass flow controllers, and a tubular reactor made from fused alumina of 11 mm o.d. The temperature was measured by means of an inert internal thermocouple and controlled by an electronic temperature controller. Product gases were analysed by the g.c. 

Conventional alternatives to a microreactor, such as non-catalytic tubular reactors, exhibit about 100-150 m m surface area to reactor volume, a value very similar to conventional heat exchangers. Using these reactors with porous catalysts filling the tubes can increase the surface area dramatically up to 10 nf m . Typically, the surface area of microreactors is in the range of lO to 10 nf m . This is the surface area only of the microreactor walls, which in general are non-porous. 

Figure 14.3 Residence time distribution in different microreactors , glued reactor without coating X, glued reactorwith coating , reactor with graphite joints. Theoretical RTD curves for tubular reactors with (solid line) Bo = 33 and (dotted line)...

In a study by Yamamoto et al., a simple microreactor was constructed by inserting an SS rod into a Pd membrane tubular reactor to investigate the effects of microcharmel size on the dehydrogenation of cyclohexane to benzene [27, 28]. As shown in Figure 10.5b, it was found that at higher temperatures, increased surface area and a longer residence time for the reactants result in greater benzene production. For a... 

Figurel2.12 Comparison ofthepolydispersityindexobtainedina multilamination microreactor (filled symbols) and in a tubular reactor (open symbols) as a function of the radial Peclet number. From Ref [12].

Most of the foundation researcdi for LRP reactions has been done using batch processes. As the development of these methods progresses towards commercialization, some scoping work has been done to investigate using continuous reactors which could offer some economic benefits. A number of these studies have been done in continuous tubular reactors which approach the size scale of microreactors. 

In Pfeifer et al. (2011) this comparison was extended for the packed bed microreactor approach, showing that, with different catalyst systems for the oxidation reaction of SO2 to SC, both approaches can be competitive with conventional tubular reactors (Table 2). 

Since the residence time varies between the channels, a tracer pulse at the inlet of the microreactor will be broadened similar to the case for a tubular reactor with axial dispersion. As a first approximation, the relative standard deviation in the residence time is twice the relative standard deviation in the channel diameter [9] ... 

It can be seen that the volumetric mass transfer coefficient, k a, is from one to two orders of magnitude larger in microreactors compared with that in conventional process equipment However, it is mainly due to a huge increase in the specific interfacial area in microreactors. The absolute value of mass transfer coefficient is of the same order of magnitude as that obtained in static mixers and in tubular reactors. 

In addition to planar microreactors (often called chips ) of the type discussed previously, capillary and tubular reactors are often used for synthetic apphcations. Clearly, as the dimension of the reactor gets bigger, the mixing efficiency itself is not as good however, by incorporation of a micromixer into the system, this problem can easily be circumvented. Furthermore, the heat transfer in larger mbes is also not as good however, for most chemical reactions, this is not a major problem as heat transfer is still several orders of magnitude better compared with batch reactors. 

Ethylene oxide catalyst research is expensive and time-consuming because of the need to break in and stabilize the catalyst before rehable data can be collected. Computer controlled tubular microreactors containing as Httle as 5 g of catalyst can be used for assessment of a catalyst s initial performance and for long-term life studies, but moving basket reactors of the Berty (77) or Carberry (78) type are much better suited to kinetic studies. 

Figure 2.29 Photograph of the ceramic reactor housing and the quartz-glass tubulare microreactor (visible through the center hole) [59].

Many speciahzed laboratory reactors and operating conditions have been used. Sinfelt has alternately passed reactants and inert materials through a tubular-flow reactor. This mode of operation is advantageous when the activity of the fixed bed of catalyst pellets changes with time. A system in which the reactants flow through a porous semiconductor catalyst, heated inductively, has been proposed for studying the kinetics of high-temperature (500 to 2000°C) reactions. An automated microreactor... 

Catalyst Evaluation. The powdered molecular sieves were evaluated following the treatment described above, without further activation. The 1-hexene isomerization and Cg aromatic isomerization tests were conducted in tubular, fixed bed, continuous flow microreactors. The catalyst bed contained one gram molecular sieve powder and one to three grams of similarly sized quartz chips used as diluent. The reactor was heated to the chosen reaction temperatures in a fluidized sand bath, and the reaction temperature was monitored by a thermocouple located m the catalyst bed. Typical runs lasted 3 to 5 hours during which samples were collected every 30 minutes. 

Alkylation of isobutane with 1-butene was carried out in a fixed bed down flow stainless steel tubular microreactor. The experiments were carried out in the gas phase at 1 atm total pressure, isobutane/l-butene molar ratio of 14 and 1-butene space velocity 1.0 h". Premixed isobutane and 1-butene (>99% purity, Matheson) was fed from a gas cylinder. A high paraffin to olefin ratio was choosen to reduce the chance of olefin dimerization. The catalysts (300 mg, 60-80 pm particle size) were activated in the reactor by calcining in air at 450°C for 4 hours. Air flow was then replaced by nitrogen and the catalyst temperature was lowered to the desired reaction temperature. The feed gas and the products were analyzed by a on-line gas chromatograph equipped with a CP-Sil PONA capillary column (length 50 m, film thickness 0.25 pm). 

---