Fixed-bed catalytic dehydrogenation of isobutane to isobutylene over chromia–alumina is carried out in adiabatic reactors in which the endothermic heat of reaction is supplied by heat stored in the catalyst bed during a preceding air-regeneration step. The reaction i-C₄H₁₀ ⇌ i-C₄H₈ + H₂ has a standard enthalpy of approximately +118–122 kJ mol⁻¹; equilibrium therefore moves toward isobutylene at low hydrocarbon partial pressure and high temperature. Published commercial data indicate a per-pass conversion of 48–62 mol% and isobutylene selectivity of 88–94 mol% when the reactor outlet pressure is held at 0.3–0.7 bar absolute and the inlet temperature is between 540–620 °C. The chromia–alumina catalyst typically contains 12–20 wt% Cr₂O₃ on a transition alumina support and is promoted with alkali metal oxides to suppress acid-catalyzed cracking and isomerization. Surface area measured by ASTM D3663-20 ranges from 30–80 m² g⁻¹, and pore volume measured by ASTM D4284-12(2017) ranges from 0.20–0.35 cm³ g⁻¹. Feed quality is a critical boundary condition: sulfur compounds must be reduced to below 10 ppmv, water below 20 ppmv, and dienes below 0.1 mol% to prevent rapid coke accumulation and irreversible chromium loss. The reaction cycle is short because coke laydown on the fixed bed suppresses active sites within minutes; commercial reactors switch through reaction, purge, regeneration, reduction, and purge steps with total cycle times of 15–40 min depending on feed composition and catalyst age. Gas composition is verified by ASTM D1945-14, and catalyst elemental composition is typically confirmed by X-ray fluorescence on fused beads.

What Operating Envelope Governs Fixed-Bed Isobutane Dehydrogenation over Chromia–Alumina?

The fixed-bed operating envelope for chromia–alumina is bounded by coking rate at low temperature, thermal cracking at high temperature, and pressure drop at high throughput. The reactor inlet temperature is controlled to 540–620 °C; below 540 °C, equilibrium conversion falls and coke precursors condense on the catalyst surface, while above 620 °C, non-catalytic thermal cracking generates methane and C₃ hydrocarbons that reduce selectivity. Liquid hourly space velocity, calculated on a liquid-volume basis at 15 °C, is maintained between 0.4 h⁻¹ and 2.0 h⁻¹; lower values increase conversion per pass but reduce volumetric productivity and accelerate coking, while higher values approach equilibrium-limited conversion and may not justify recycle energy. The reactor outlet pressure is kept between 0.3 bar and 0.8 bar absolute to shift equilibrium toward products; the vacuum or compressor load is balanced against downstream separation pressure. If steam is used as a diluent and heat carrier, the steam-to-isobutane molar ratio is held between 0.5 and 1.5, although some commercial fixed-bed chromia–alumina variants operate without steam and rely solely on bed heat capacity. Regeneration air inlet temperature is set at 600–750 °C; the bed must not exceed 800 °C to avoid converting Cr₂O₃ to volatile CrO₃ and sintering the alumina support. Table 1 summarizes the operating envelope and the equipment or standard used for verification.

Catalyst deactivation in fixed-bed chromia–alumina service occurs through coking, chromium redox imbalances, and hydrothermal sintering; these mechanisms operate on different timescales and require different regeneration responses. Coke deposition is the fastest deactivation route, with carbonaceous overlayers blocking micropores and coordinatively unsaturated Cr(III) sites. Post-reaction coke levels of 0.2–1.5 wt% are typical, and the coke is not uniformly distributed: the inlet zone of the bed accumulates higher carbon due to lower temperature and higher hydrocarbon partial pressure. Chromium redox cycling contributes to irreversible deactivation when Cr(VI) formed during air regeneration is not fully re-reduced by the next hydrocarbon cycle; residual Cr(VI) can migrate along the alumina surface and collapse isolated Cr(III) active ensembles. Regeneration air must be stepped from low oxygen concentration to 2–4 mol% oxygen at the outlet while maintaining the bed below 800 °C to limit CrO₃ volatilization. Hydrothermal sintering of the alumina support becomes significant when steam partial pressure exceeds 0.5 bar absolute and regeneration temperatures exceed 750 °C; this reduces BET surface area by 20–40% over a campaign and increases the fraction of α-Al₂O₃. Alkali promoters such as K₂O or Na₂O at 0.5–2.5 wt% neutralize strong acid sites but also increase the mobility of chromium species under hydrothermal conditions. Table 2 lists deactivation modes and their monitoring methods.

When catalyst regeneration air rates fall below carbon burn-off stoichiometry

Complete combustion of deposited coke to carbon dioxide requires 11.5 kg air per kilogram of carbon, based on 21 mol% oxygen in air and a molecular weight of 29 kg kmol⁻¹. If the regeneration air flow is insufficient, partial oxidation to carbon monoxide dominates; the heat release drops from approximately 32.8 MJ kg⁻¹ carbon for CO₂ formation to 9.2 MJ kg⁻¹ carbon for CO formation. The fixed bed then fails to reach the target regeneration temperature, and the next dehydrogenation cycle starts with a lower bed temperature and a shorter equilibrium-limited run length. Commercial regeneration control therefore maintains excess oxygen at 2–4 mol% dry at the regeneration effluent and limits carbon monoxide breakthrough to below 1 mol% using an infrared CO analyzer. The air rate must also be corrected for residual hydrocarbons in the bed after the purge step; without complete purge, combustible gas mixed with regeneration air can generate local hot spots above 800 °C and sinter the chromia–alumina support. The regeneration sequence is normally divided into a low-oxygen ignition phase, a main burn-off phase, and a high-temperature heat-soak phase. During the main burn-off phase, the bed temperature rise is typically limited to 50–80 °C across the bed by adjusting air flow and nitrogen dilution. Failure to maintain air flow also leaves carbon in the catalyst pores; this residual carbon reduces the effective Cr(III) active-site density at the start of the next reaction step and increases the rate of pressure-drop build-up in the reactor. Regeneration blower discharge pressure must exceed the bed pressure drop plus the control valve margin; if the blower operates below its rated discharge, air flow is corrected by the control system, but a low-flow alarm below 70% of design air flow should be interlocked to stop hydrocarbon feed and prevent endothermic cooling of an only partially regenerated bed.

Fixed-bed pressure drop and pellet crush-strength constraints

Pressure drop across a fixed-bed dehydrogenation reactor is governed by the Ergun equation: ΔP/L = 150 [(1−ε)² μ u_s / (ε³ d_p² φ²)] + 1.75 [(1−ε) ρ u_s² / (ε³ d_p φ)], where ε is bed void fraction, d_p is pellet diameter, φ is sphericity, u_s is superficial gas velocity, μ is gas viscosity, and ρ is gas density. For a 4 mm extrudate with bed void fraction 0.36–0.40, sphericity 0.75–0.85, and superficial gas velocity 0.5–1.0 m s⁻¹ at 550 °C, pressure drop is typically 10–25 kPa m⁻¹ bed height. The maximum allowable bed pressure drop in commercial fixed-bed dehydrogenation is frequently set at 50 kPa to prevent particle breakage and gas channelling. Single-pellet crush strength measured by ASTM D4179-22 must remain above 25 N; lower values lead to fines generation during thermal cycling, and fines fill the void spaces, increasing pressure drop and causing flow maldistribution. The catalyst loading procedure must exclude particles smaller than 1 mm by sieving, and the reactor internals should include a distribution plate and screen to prevent channeling. Because chromia–alumina catalysts are dense and relatively brittle, the bed is heated and cooled at rates below 50 °C h⁻¹ during start-up and shutdown to avoid thermal shock fracture. In radial-flow fixed beds, the catalyst is contained in annular baskets with center-pipe collection, and the radial pressure drop is designed to be 5–15 kPa to maintain uniform flow across the bed. Reactor vessels are typically carbon steel with refractory lining; the refractory must withstand the regeneration temperature and the reducing atmosphere of the hydrocarbon cycle without silicate migration.

Downstream of the fixed-bed reactor, the effluent is quenched to stop secondary cracking and prevent the reverse hydrogenation of isobutylene. The gas mixture contains unconverted isobutane, isobutylene, hydrogen, methane, ethane, propylene, and minor C₄ diolefins; water cooled to 30–40 °C removes heavy hydrocarbons and a portion of the steam if steam dilution is used. The C₄ fraction is then compressed to 10–16 bar gauge and sent to a selective hydrogenation unit where palladium/alumina converts acetylenic and diene impurities to monoolefins at 40–80 °C; acetylene content is reduced to below 5 ppmv and total dienes below 50 ppmv. The isobutylene product is separated from unconverted isobutane by extractive distillation with a solvent such as N-methylpyrrolidone or by using a C₄ splitter after MTBE etherification; the choice depends on whether the plant sends isobutylene to methyl tert-butyl ether production or to high-purity isobutylene recovery. For MTBE synthesis over sulfonic acid resin, the feed must meet a maximum sulfur content of 1 ppmw, water below 100 ppmw, and a bromine index below 50 mg Br/100 g according to ASTM D1159-07(2017). If high-purity isobutylene is required, residual n-butane and 1-butene are removed by a combination of superfractionation and adsorption; this configuration increases capital cost and energy consumption but raises isobutylene purity above 99.0 wt%. Spent chromia–alumina catalyst and chromium-containing dust from the reactor are classified as hazardous waste when leachable chromium exceeds 5 mg L⁻¹ in the toxicity characteristic leaching procedure of US EPA SW-846 Method 1311. In the European Union, chromium(VI) compounds are classified as Carcinogen Category 1B and Mutagen Category 1B under CLP Regulation (EC) No 1272/2008, and handling must follow the occupational exposure limits in Directive 2004/37/EC. Spent catalyst passivation should therefore be conducted with dilute hydrogen at 350–400 °C until the Cr(VI) fraction is below 0.1 wt%; the reactor is then cooled under nitrogen and discharged under high-integrity dust containment.