Four-lump kinetic model for fluid catalytic cracking process

In the fluid catalytic cracking reactor heavy gas oil is cracked into more valuable lighter hydrocarbon products. The reactor input is a mixture of hydrocarbons which makes the reaction kinetics very complicated due to the involved reactions. In this paper, a four-lump model is proposed to describe the process. This model is different from others mainly in that the deposition rate of coke on catalyst can be predicted from gas oil conversion and isolated from the C₁–C₄ gas yield. This is important since coke supplies heat required for endothermic reactions occurring in the reactor. By this model we can also conclude that the C1–C4 gas yield increases with increasing reactor temperature, while production of gasoline and coke decreases.     

Depolymerization of clean unfilled PTFE waste in a continuous process

A continuous process is described whereby waste PTFE was converted into tetrafluoroethylene (C₂F₄, TFE), hexafluoroethane (C₂F₆, HFE), hexafluoropropylene (C₃F₆, HFP), and octafluorocyclobutane (c‐C₄F₈, OFCB). The waste PTFE was depolymerized inside a reactor that was heated by a radiofrequency induction generator. The reactor was capable of operating at various temperatures (600–900°C) as well as various reduced pressures (5–80 kPa). The depolymerization reaction conditions could be changed while the reactor was in operation in order to manipulate the reaction product composition. Under certain conditions, high yields of TFE (> 94%) and low concentrations of by‐products were formed. The PTFE was fed vertically downward from the hopper, with the screw feeder, into the reactor where the depolymerization process took place. The hot intermediate products were continuously evacuated through a self‐cleaning quench probe, where it was quenched to form the final products. The depolymerized products were analyzed with a gas chromatograph. © 2006 Wiley Periodicals, Inc. J Appl Polym Sci 102: 2752–2759, 2006     

The kinetics modeling and reactor simulation of propylene chlorination reaction process

The kinetics experiments of fast reaction process of propylene chlorination at low temperature (30–90°C) and high temperature (420–480°C) are respectively conducted, and the corresponding reaction mechanisms and kinetics models are proposed. The radical mechanism at high temperature and the molecular mechanism at low temperature are found to be most likely with the experimental results. Specifically, the kinetics model, firstly considering the reversible reaction step of forming C₃H₆Cl · and direct hydrogen abstraction of forming C₃H₅ · , shows better agreement with the experimental data. Furthermore, the critical reaction temperature T_critical is firstly proposed to determine the dominant reaction mechanism in different conditions, and correspondingly the combination method of the high-temperature and low-temperature kinetics models has been adopted for tubular reactor simulation, which can reasonably reflect the influence of wide variation range of temperature in the reactor and guide the industrial reactor design in the further work.     

Differential scanning calorimetry studies of the cure of carbon fiber–epoxy composite prepregs

Diaminodiphenyl sulfone (DDS) cured tetraglycidyl-4,4'-diaminodiphenyl methane (TGDDM) epoxies, whose cure reactions are accelerated by BF₃:amine catalysts, are the most common composite matrices utilized in aerospace high performance, fibrous composites. To process reproducible composites requires an understanding of the cure reactions and how these reactions are modified by the BF₃:amine catalysts. In this article we report systematic differential scanning calorimetry (DSC) studies of (i) the constituents of BF₃:NH₂C₂H₅-catalyzed TGDDM–DDS epoxies and their mixtures, (ii) the effect of BF₃:NH₂C₂H₅ concentration on the cure reactions, (iii) the nature of the catalyzed cure reactions, and (iv) the environmental sensitivity of the catalyst. DSC studies are also reported on the cure reaction characteristics and environmental sensitivity of commercial C fiber–TGDDM–DDS epoxy prepregs.     

Pyrolysis of used sunflower oil in the presence of sodium carbonate by using fractionating pyrolysis reactor

Pyrolysis of used sunflower oil was carried out in a reactor equipped with a fractionating packed column (in three different lengths of 180, 360 and 540 mm) at 400 and 420°C in the presence of sodium carbonate (1, 5, 10 and 20% based on oil weight) as a catalyst. The use of packed column increased the residence times of the primer pyrolysis products in the reactor and packed column by the fractionating of the products which caused the additional catalytic and thermal reactions in the reaction system and increased the content of liquid hydrocarbons in gasoline boiling range. The conversion of oil was high (42–83 wt.%) and the product distribution was depended strongly on the reaction temperature, packed column length and catalyst content. The pyrolysis products consisted of gas and liquid hydrocarbons, carboxylic acids, CO, CO₂, H₂ and water. Increase in the column length increased the amount of gas and coke–residual oil and decreased the amount of liquid hydrocarbon and acid phase. Also, increase of sodium carbonate content and the temperature increased the formation of liquid hydrocarbon and gas products and decreased the formation of aqueous phase, acid phase and coke–residual oil. The major hydrocarbons of the liquid hydrocarbon phase were C₅–C₁₁ hydrocarbons. The highest C₅–C₁₁ yields (36.4%) was obtained by using 10% Na₂CO₃ and a packed column of 180 mm at 420°C. The gas products included mostly C₁–C₃ hydrocarbons.     

Olefin production by cofeeding methanol and n‐butane: Kinetic modeling considering the deactivation of HZSM‐5 zeolite

The joint transformation of methanol and n‐butane fed into a fixed‐bed reactor on a HZSM‐5 zeolite catalyst has been studied under energy neutral conditions (methanol/n‐butane molar ratio of 3/1). The kinetic scheme of lumps proposed integrates the reaction steps corresponding to the individual reactions (cracking of n‐butane and MTO process at high‐temperature) and takes into account the synergies between the steps of both reactions. The deactivation by coke deposition has been quantified by an expression dependent on the concentration of the components in the reaction medium, which is evidence that oxygenates are the main coke precursors. The concentration of the components in the reaction medium (methanol, dimethyl ether, n‐butane, C₂? C₄ paraffins, C₂? C₄ olefins, C₅? C₁₀ lump, and methane) is satisfactorily calculated in a wide range of conditions (between 400 and 550°C, up to 9.5 (g of catalyst) h (mol CH₂)^?1 and with a time on stream of 5 h) by combining the equation of deactivation with the kinetic model of the main integrated process. © 2010 American Institute of Chemical Engineers AIChE J, 2011     

Fischer–Tropsch synthesis in slurry-phase reactors over Mn- and Zr-modified Co/SiO2 catalysts

Fischer–Tropsch (F–T) synthesis was carried out in a gas-flowed slurry-phase reaction system over Mn- and Zr-modified Co/SiO₂ catalysts. A 0.5 L stirred tank slurry reactor (STSR) was used for catalyst screening and a 12.5 L slurry bubble column reactor (SBCR) was used for trial pilot operation. While using the 0.5 L reactor for catalyst screening, Co supported on the SiO₂ with an average pore size of 10 nm showed a high catalytic performance for the F–T synthesis due to the suitable Co particle size in the catalyst. Zr promoter improved the activity and Mn promoter improved the stability of Co/SiO₂ catalyst for the F–T synthesis. H₂-TPR profiles indicated that Zr and Mn promoters improved the reduction degree of Co₃O₄ particles (on SiO₂ surface) to Co⁰ active species in H₂ flow at low temperature. While using the 12.5 L reactor for trial pilot operation over Mn–Zr–Co/SiO₂ catalyst, the space-time yield (STY) of C₅₊ hydrocarbons (liquid fuel) showed almost the same values when various solvents (n-C₁₆H₃₄, n-C₁₄H₃₀, diesel from petrol station, F–T crude oil) were used. Diesel and F–T crude oil are suitable for using in a large-scaled F–T synthesis plant owing to the low prices. Mn–Zr–Co/SiO₂ catalyst achieved a STY of C₅₊ hydrocarbons larger than 1000 g-C₅₊ kg-cat^? 1 h^? 1 in the 12.5 L reactor. The production capacity of liquid fuel from the 12.5 L reactor reached to 15.6 L per day (assumed for 24 h continuous operation). The stirring was very important for the F–T synthesis both reaction in the 0.5 L reactor and reaction in the 12.5 L reactor. The shape of slurry reactor also influenced the CO conversion for the F–T synthesis: reaction in the 12.5 L SBCR gave a higher CO conversion than that of reaction in the 0.5 L STSR (at the same W/F value under the same stirring speed) because the slender column reactor (SBCR) extended the residue time of reaction gas in the slurry-phase containing catalyst.     

Hydrotreatment of jatropha oil over noble metal catalysts

Jatropha oil is a promising nonedible feedstock for producing renewable diesel. In this work, the hydrotreatment processing of jatropha oil was investigated. Instead of using conventional alumina-supported Co–Mo, Ni–Mo, and Ni–W catalysts that need sulfidation pretreatment, noble metals such as Pd and Ru were chosen. Trials were performed in an isothermal trickle-bed reactor and the reaction conditions were as follows: temperature 603–663?K, weight hourly space velocity (WHSV) 1 to 4/h, pressure 1.5–3?MPa, and H₂/oil ratio 200–800 (v/v). Yield of n-C₁₅ to n-C₁₈ hydrocarbons was maximized (70.3 and 43.8% for Pd/Al₂O₃ and Ru/Al₂O₃, respectively) at the following conditions: T?=?663 K, WHSV?=?2/h, P?=?3?MPa, and H₂/oil ratio?=?600 (v/v). Since Ru favored cracking reactions to a larger extent than Pd, the yield of C₁₅ to C₁₈ hydrocarbons over Ru/Al₂O₃ was lowered. Using simple first-order plots for oil conversion, activation energies for the hydrotreating process over Pd/Al₂O₃ and Ru/Al₂O₃ were found and they were equal to 109 and 121?kJ/mol, correspondingly.     

Modelling of the gas phase chemistry during diamond CVD: the role of different hydrocarbon species

The chemkin suite of computer programs has been used to model the concentration profiles of different hydrocarbon species present within a hot filament CVD reactor during diamond growth, and the calculated values are compared with those obtained by direct measurements using an in situ molecular beam mass spectrometer. Different hydrocarbon gases (CH₄, C₂H₂, C₂H₄ and C₂H₆) were used as the carbon source in the input gas mixture, ensuring that the ratio of C:H₂ remained constant at 1%. Calculations for when C₂H₆ is used as the precursor gas, after reaction and thermal equilibrium is realised, yield similar CH₄:C₂H₂ mole fraction ratios in the reactor under growth conditions to those obtained using CH₄, and to those measured experimentally. However, simulations using C₂H₄ or C₂H₂ as input gases do not reproduce the experimentally observed ratio of CH₄:C₂H₂ mole fractions. This suggests that the conversion of unsaturated C₂ species to C₁ species is not a straightforward gas phase process, and there must be one or more reactions occurring within the chamber that are not present in the standard models for hydrocarbon reactions. We suggest that these neglected reaction(s) probably involve surface-catalysed hydrogenation, which in this case, is most likely occurring on the surface of the filament.     

A novel approach for describing mixing effects in homogeneous reactors

The Liapunov–Schmidt technique of classical bifurcation theory is used to spatially average the convection–diffusion–reaction (CDR) equations over smaller time/length scales to obtain low-dimensional two-mode models for describing mixing effects due to local diffusion, velocity gradients and reactions. For the cases of isothermal homogeneous tubular, loop/recycle and tank reactors, the two-mode models are described by a pair of coupled balance equations for the mixing-cup (C_m) and spatial average (〈C〉) concentrations. The global equation describes the variation of C_m with residence time (or position) in the reactor, while the local equation expresses the coupling between local diffusion, velocity gradients and reaction at the local scales, in terms of the difference between C_m and 〈C〉. It is shown that the two-mode models have many similarities with the classical two-phase models of heterogeneous catalytic reactors with the concept of transfer between phases being replaced by that of exchange between the two-modes. It is also shown that when the local Damköhler number (ratio of local diffusion to reaction time) is small, the solution of two-mode models approaches the exact solution of full CDR equations, while for fast reactions the two-mode models retain all the qualitative features of the latter. Examples are provided to illustrate the usefulness of these two-mode models in predicting micromixing effects on homogeneous reactions.     

Thermal decomposition of trichloroethylene under a reducing atmosphere of hydrogen

The thermal reaction of trichloroethylene (TCE: C₂HCl₃) has been conducted in an isothermal tubular flow reactor at 1 atm total pressure in order to investigate characteristics of chlorinated hydrocarbons decomposition and pyrolytic reaction pathways for formation of product under excess hydrogen reaction environment. The reactions were studied over the temperature range 650 to 900 °C with reaction times of 0.3–2.0 s. A constant feed molar ratio C₂HCl₃: H₂ of 4: 96 was maintained through the whole experiments. Complete decay (99%) of the parent reagent, C₂HCl₃ was observed at temperature near 800 °C with 1 s reaction time. The maximum concentration (28%) of C₂H₂Cl₂ as the primary intermediate product was found at temperature 700 °C where up to 68% decay of C₂HCl₃ occurred. The C₂H₃Cl as highest concentration (19%) of secondary products was detected at 750 °C. The one less chlorinated methane than parent increased with temperature rise subsequently. The number of qualitative and qualitative chlorinated products decreased with increasing temperature. HCl and dechlorinated hydrocarbons such as C₂H₄, C₂H₆, CH₄ and C₂H₂ were the final products at above 800 °C. The almost 95% carbon material balance was given over a wide range of temperatures, and trace amounts of C₆H₆, C₄H₆ and C₂HCl were observed above 800 °C. The decay of reactant, C₂HCl₃ and the hydrodechlorination of intermediate products, resulted from H atom cyclic chain reaction via abstraction and addition replacement reactions. The important pyrolytic reaction pathways to describe the important features of reagent decay, intermediate product distributions and carbon mass balances, based upon thermochemical and kinetic principles, were suggested. The main reaction pathways for formation of major products along with preliminary activation energies and rate constants were given.     

The cure reactions,network structure,and mechanical response of diaminodiphenyl sulfone-cured tetraglycidyl 4,4′diaminodiphenyl methane epoxies

The cure reactions, chemical structure, and network topography of diaminodiphenyl sulfone (DDS)-cured tetraglycidyl 4,4′diaminodiphenyl methane (TGDDM) epoxies are reported. Systematic Fourier transform infrared (FTIR) spectroscopy studies of the cure and degradation reactions of TGDDM-DDS epoxies in the 100–300°C temperature range as a function of DDS and boron trifluoride monoethylamine, BF₃: NH₂C₂H₅ catalyst concentrations are presented. FTIR studies indicate TGDDM epoxide homopolymerizes in the 175–250°C range via epoxide-hydroxyl (E-OH) chain extension reactions. The hydroxyl groups are initially present as α-glycol impurities or are formed by epoxide isomerization and/or oxidation. Three principal cure reactions occur at 177°C for TGDDM-DDS epoxies, namely primary amine-epoxide (PA-E), secondary amine-epoxide (SA-E) and E-OH reaction with the PA-E reaction being an order of magnitude faster than the other two cure reactions. The PA-E reaction dominates the early stages of cure and, hence, composite processing conditions. The three cure reactions are catalyzed to similar extents by BF₃: NH₂C₂H₅. FTIR and molecular modeling studies indicate that the E-OH and SA-E reactions can occur intermolecularly to form crosslinks or intramolecularly to form non-cross-linked internal rings. The complex degradation reactions of TGDDM-DDS epoxies in the 177–300°C range are reported. The principal degradation reactions involve (1) dehydration and/or oxidation to form ether crosslinks and (2) decomposition of the EOH cure reaction products to form propenal. Based on a knowledge of the cure reactions, together with molecular modeling, the chemical structure and network topography of TGDDM-DDS epoxies are reported.     

Enhanced Production of C2–C4 Olefins Directly from Synthesis Gas

An attempt made for the selective production of C₂–C₄ olefins directly from the synthesis gas (CO + H₂) has led to the development of a dual catalyst system having a Fischer–Tropsch (K/Fe–Cu/AlOx) catalyst and cracking (H-ZSM-5) catalyst operate in consecutive dual reactors. The flow rate (space velocity) and H₂/CO molar ratio of the feed have been optimized for achieving higher CO conversions and olefin selectivities. The selectivity to C₂–C₄ olefins is further enhanced by optimizing the reaction temperature in the second reactor (cracking), where the product exhibited 51% selectivity to C₂–C₄ hydrocarbons rich in olefins (77%) with a stable time-on-stream performance in a studied period of 100 h.     

HC-SCR reaction pathways on ion exchanged ZSM-5 catalysts

Metal cation (metal = Cu, In and La) ion exchanged ZSM-5 zeolites as catalysts for the NO selective reduction by propane and propene in excess oxygen. The surface reactions of HC-SCR over catalysts were investigated through in situ DRIFTS method. For C₃H₈-SCR, adsorbed nitrate species (–NO₃) were observed as main reaction intermediates and they could react with gaseous propane to produce N₂, H₂O and CO₂. While for C₃H₆-SCR, adsorbed amine species (–NH₂) were observed as main reaction intermediates and they could react with NO or NO₂ to produce the final products. The different reaction pathways for C₃H₈-SCR and C₃H₆-SCR over catalysts were proposed based on the DRIFTS results and the main factors controlling the activities of catalysts were discussed in details. The competing adsorption between NO–O₂ and HC–O₂ on the Brønsted acid sites of catalysts was responsible for the different reaction pathways in HC-SCR.     

Reactor performance and stability in an alternating reaction-reheat paraffin dehydrogenation system

The classical fixed bed C₃–C₄ paraffin dehydrogenation process is a cyclic operation in which the reactor alternates between reaction and reheat cycles. During the reheat cycle, the necessary energy for the dehydrogenation reaction is stored in the fixed bed by passing hot air through it. In this established technology, both the hydrocarbon reactant and the reheat hot air are fed into the fixed bed from the same end (top) of the reactor. This is termed parallel flow (cocurrent) operation. An alternative feeding fixed bed has the hydrocarbon reactant and the reheat air entering from the opposite ends of the reactor. This is termed reverse flow (countercurrent) operation. This alternate creates an ideal temperature profile for an equilibrium limited endothermic reaction (rising temperature profile along the reactor). The transient flow behavior of both parallel and reverse flow reactors has been modelled and the dynamics of temperature profile development for both concepts have been analyzed. Based upon the model predictions, the characteristics as well as the reactor stability of the both concepts have been discussed.     

Enzymatic synthesis of medium‐chain triacylglycerols by alcoholysis and interesterification of copra oil using a crude papain lipase preparation

We have examined the possibility of producing analogs of medium‐chain triglycerides (MCT) from copra oil, i.e. a triacylglycerol mixture with a high content of medium‐chain fatty acid moieties (C₆–C₁₀). A two‐step enzymatic process was used in which copra triacylglycerols were first split with papain lipase by alcoholysis with an alkyl alcohol and then subjected to interesterification with the alkyl esters recovered using papain lipase. Effects of temperature, water activity content, substrate ratio, biocatalyst amount, and alcohol chain length were also investigated. On the one hand, the sn‐3 stereoselectivity of the lipase in the alcoholysis of copra oil with butanol has permitted a direct enrichment of caproic, caprylic and capric moieties in the synthesized butyl esters. Thus, in the batch reactor, the reaction led to about 31% conversion of the oil after 24 h, and the content of C₆–C₁₀ acids in the synthesized esters increased from about 16% in the starting oil to almost 42%. A similar enzymatic alcoholysis in a packed‐bed column bioreactor gave 31% conversion of the oil after 120 min of reactor residence time. The reaction was also very selective because the C₆–C₁₀ fatty acyl groups represented about half of the newly formed butyl esters, whereas they accounted for only 16% of total fatty acids in the starting oil. On the other hand, the transesterification of the alkyl esters recovered (highly enriched in C₆–C₁₀ fatty acyl groups) with native copra oil directly led to an increase in the content of MCT in the oil, from 18 mol‐% at the beginning of the reaction to 61 mol‐% of MCT after a time period of 72 h in the batch reactor.     

Epoxidation of a Series of C2–C6 Olefins in the Gas Phase

Epoxidation of ethylene, propylene, 2‐methylpropene, trans‐2‐butene, 2‐methyl‐2‐butene, and 2,3‐dimethyl‐2‐butene were carried out in a flow‐through reactor in the homogeneous gas phase at pressures of 0.25–1.0 bar in the temperature range of 250–375 °C. Residence times in the reactor varied from 8.3 to 38 ms. The oxidizing agent needed in the feed gas is ozone. The O₃ efficiency (reacted olefin/initial O₃) was found to be strongly dependent on the reactivity of the olefin used. For C₄–C₆ olefins, the O₃ efficiency was better than 75 % in each case. For 2‐methyl‐2‐butene and 2,3‐dimethyl‐2‐butene, the O₃ efficiency exceeded the theoretical value of 100 % considerably. The selectivity to epoxide was about 90 % independent of the olefin used. Under conditions of nearly total olefin conversion, the high selectivity to the epoxide has been retained as unchanged. There were no indications for consecutive reactions of the epoxides.     

Assessing the performance of an industrial SBCR for Fischer–Tropsch synthesis: Experimental and modeling

The main objective of this study is to predict the performance of an industrial‐scale (ID = 5.8 m) slurry bubble column reactor (SBCR) operating with iron‐based catalyst for Fischer–Tropsch (FT) synthesis, with emphasis on catalyst deactivation. To achieve this objective, a comprehensive reactor model, incorporating the hydrodynamic and mass‐transfer parameters (gas holdup, ε_G, Sauter‐mean diameter of gas bubbles, d₃₂, and volumetric liquid‐side mass‐transfer coefficients, k_La), and FT as well as water gas shift reaction kinetics, was developed. The hydrodynamic and mass‐transfer parameters for He/N₂ gaseous mixtures, as surrogates for H₂/CO, were obtained in an actual molten FT reactor wax produced from the same reactor. The data were measured in a pilot‐scale (0.29 m) SBCR under different pressures (4–31 bar), temperatures (380–500 K), superficial gas velocities (0.1–0.3 m/s), and iron‐based catalyst concentrations (0–45 wt %). The data were modeled and predictive correlations were incorporated into the reactor model. The reactor model was then used to study the effects of catalyst concentration and reactor length‐to‐diameter ratio (L/D) on the water partial pressure, which is mainly responsible for iron catalyst deactivation, the H₂ and CO conversions and the C₅₊ product yields. The modeling results of the industrial SBCR investigated in this study showed that (1) the water partial pressure should be maintained under 3 bars to minimize deactivation of the iron‐based catalyst used; (2) the catalyst concentration has much more impact on the gas holdup and reactor performance than the reactor height; and (3) the reactor should be operated in the kinetically controlled regime with an L/D of 4.48 and a catalyst concentration of 22 wt % to maximize C₅₊ products yield, while minimizing the iron catalyst deactivation. Under such conditions, the H₂ and CO conversions were 49.4% and 69.3%, respectively, and the C₅₊ products yield was 435.6 ton/day. © 2015 American Institute of Chemical Engineers AIChE J, 61: 3838–3857, 2015     

Brown coal conversion under the action of supercritical water

A test bench was developed and the conversion of the organic matter of coal (OMC) in supercritical water (SCW) was studied under conditions of a continuous supply of a water-coal suspension to a vertical flow reactor at 390–760°C and a pressure of 30 MPa. From 44 to 63% OMC was released as liquid and gaseous products from coal particles (from the water-coal supension) during the time of fall to the reactor. This stage was referred to as the dynamic conversion of coal. The particles passed through the stage of the dynamic conversion of coal did not agglomerate in the reactor in the subsequent process of batch conversion in a coal layer at T = 550–760°C. The volatile products of the overall process of the dynamic and batch conversion of coal included saturated hydrocarbons (CH₄ and C₂H₆), aromatic hydrocarbons (C₆H₆, C₇H₈, and C₈H₁₀), synthesis gas (H₂ and CO), and CO₂. At T < 600°C, CO₂ and CO were the degradation products of oxygen-containing OMC fragments, whereas they also resulted from the decomposition of water molecules at higher temperatures in accordance with the reaction (C) + H₂O = CO + H₂. The mechanisms were considered, and the parameters responsible for the dynamic conversion of coal were calculated.     

Synthesis of light olefins from syngas over Fe–Mn–V–K catalysts in the slurry phase

Fe–Mn–V–K catalysts for the synthesis of light olefins from CO hydrogenation were prepared by a specially controlled degradation method. The effect of the V content on the structure and the catalytic performance of the catalysts were investigated in a continuously stirred tank slurry reactor. Mössbauer spectra (MES) results show that the incorporation of V with appropriate contents can improve the dispersion of the α-Fe₂O₃ phase. CO hydrogenation results indicate that a small addition of V can improve the product distribution. The addition can also increase the selectivity to light olefins by inhibiting the secondary hydrogenation reaction of the initial olefin. The best catalytic performance was obtained at the Fe/Mn/V molar ratio of 3/1/0.2. The total C₂–C₄ content in all hydrocarbons and O/P in the C₂–C₄ fraction were 49.15 wt% and 3.95, respectively.