Hydrodealkylation of heavy aromatic streams containing alkylnaphthalenes, C₁₁–C₁₃ homologues, and pre-existing naphthalene is carried out in fixed-bed adiabatic reactors where feed naphthalene content above 8 wt% alters the molar selectivity to naphthalene from mono-methylnaphthalene conversion by superimposing equilibrium hydrogenation losses onto the dealkylation pathway. Plant-derived mass balances on naphthalene production units indicate that a feed naphthalene concentration above 8 wt% frequently corresponds to a heavy reformate or light cycle oil cut with an initial boiling point above 210°C and a final boiling point below 300°C, as determined by ASTM D2887 simulated distillation. In this regime, the feed already contains significant condensed-ring aromaticity; the incremental naphthalene produced from 2-methylnaphthalene and 1-methylnaphthalene must be distinguished from the inlet naphthalene via component-level around-reactor gas chromatography, typically following UOP 744 or an equivalent internal standard method. Published data for the exact selectivity shift at feed naphthalene contents between 8 wt% and 15 wt% remain limited because most licensor studies normalise the feed to 6–8 wt% naphthalene to avoid precipitation in the feed train and because analytical closure for condensed-ring aromatics is difficult below 0.1 wt% detection limits. Nevertheless, two mechanisms are consistently reported in hydrodealkylation literature: the higher naphthalene partial pressure increases tetralin formation through a reversible hydrogenation route, and the higher condensed-ring concentration increases the fouling rate in the preheat exchanger and first catalyst bed.

Thermal degradation thresholds in naphthalene-rich hydrodealkylation feeds

The reactor inlet temperature in catalytic hydrodealkylation of naphthalene-rich feeds is normally maintained between 580°C and 700°C, with hydrogen partial pressure between 20 bar and 50 bar. At these temperatures, thermal dealkylation of alkylnaphthalenes proceeds by homolytic cleavage of the aryl–methyl bond, while catalytic hydrogenation of naphthalene to tetralin becomes increasingly reversible with rising temperature; the standard Gibbs free energy of naphthalene hydrogenation changes sign within this temperature window, moving the thermodynamic equilibrium toward naphthalene but not eliminating kinetically accessible tetralin at the reactor outlet. When the feed naphthalene content exceeds 8 wt%, the higher concentration of condensed-ring aromatic compounds increases the heat release per unit mass of liquid feed because naphthalene hydrogenation to tetralin is exothermic by approximately −125 kJ/mol; this exotherm can generate a temperature rise of 30–60°C across an adiabatic bed, depending on hydrogen-to-hydrocarbon ratio and liquid hourly space velocity. The resulting axial temperature profile must be managed with quench hydrogen injection, because a local temperature excursion above 720°C initiates rapid carbon deposition through condensed-ring condensation reactions, as documented in fixed-bed pilot-plant studies. Preheat train fouling is also a processing hazard at naphthalene contents above 8 wt%; naphthalene solidifies at 80.2°C, and high-boiling aromatic condensation products can raise the viscosity of the liquid feed, necessitating steam tracing of transfer lines and thermal insulation of the charge pump suction. Analytical monitoring of the feed by ASTM D2425 hydrocarbon type analysis and of the reactor effluent by gas chromatography–mass spectrometry provides the quantitative basis for distinguishing dealkylation selectivity from hydrogenation losses, but the uncertainty in reported selectivity increases when the feed naphthalene-to-methylnaphthalene ratio exceeds 1.5.

Because the feed naphthalene content influences the apparent selectivity calculation, units that do not use internal-standard component balances may report depressed selectivity values at feed naphthalene concentrations above 8 wt% even when the intrinsic dealkylation selectivity from methylnaphthalene to naphthalene remains constant. In a single-pass reactor with a naphthalene recovery column, the measured net naphthalene selectivity is the difference between naphthalene produced from alkylnaphthalene dealkylation and naphthalene consumed by hydrogenation to tetralin or ring-opening to benzene, toluene, and lighter hydrocarbons. For a feed containing 10 wt% naphthalene, a tetralin yield of 2–4 wt% based on feed naphthalene can reduce the apparent selectivity by 5–10 percentage points, even though the methylnaphthalene conversion remains above 90%. This analytical artefact is particularly pronounced in older units where the feed naphthalene concentration is calculated from a naphthalene recovery mass balance rather than from direct gas chromatographic analysis of the reactor feed. To resolve the intrinsic selectivity, cracked-gas and aromatic product streams are analysed for C₁₀–C₁₂ aromatic distribution by ASTM D5769 or equivalent gas chromatography–mass spectrometry, and hydrogen consumption is metered with Coriolis or orifice-plate mass flow devices traceable to ISO 5167. The resulting component flow rates allow calculation of the molar selectivity as the molar flow of naphthalene in the reactor outlet attributable to methylnaphthalene conversion divided by the molar flow of methylnaphthalene converted, but the calculation requires accurate accounting for naphthalene entering and leaving the reactor. Published data for this specific configuration are limited, but plant mass balances show that the apparent selectivity correction becomes necessary when the feed naphthalene content exceeds 8 wt% because the signal-to-noise ratio of incremental naphthalene product decreases as the feed concentration rises.

How Does Excess Naphthalene Alter Catalytic Site Occupancy?

Naphthalene and methylnaphthalenes adsorb strongly on sulfided nickel–molybdenum and cobalt–molybdenum catalysts, but naphthalene adsorbs preferentially on the hydrogenation sites responsible for aromatic ring saturation, whereas methylnaphthalenes require metal–acid bifunctional sites for alkyl transfer and dealkylation. At feed naphthalene concentrations above 8 wt%, the surface concentration of naphthalene on the hydrogenation sites is increased, reducing the availability of activated hydrogen for the demethylation of adsorbed methylnaphthalene intermediates. This competitive adsorption effect is consistent with observed decreases in hydrodealkylation selectivity at hydrogen-to-hydrocarbon molar ratios below 4; at hydrogen-to-hydrocarbon ratios above 6, the selectivity difference between feeds with 6 wt% and 12 wt% naphthalene narrows, indicating that hydrogen availability rather than thermodynamic equilibrium controls the product distribution. The catalyst life in a naphthalene-rich feed is also influenced by the rate of polycondensation of adsorbed naphthalene and tetralin intermediates to heavier triaromatic and tetraaromatic species, which can block micropores and reduce the effective diffusivity of methylnaphthalenes. In fixed-bed reactors with a catalyst particle diameter of 1.6–3.0 mm, intraparticle diffusion limitations become significant at temperatures above 600°C, and the Thiele modulus for methylnaphthalene conversion increases as the feed naphthalene concentration rises because the heavier aromatic compounds occupy pore volume. Regeneration intervals in naphthalene-rich service are therefore often shorter than in operations where the feed naphthalene content is maintained below 8 wt% by blending with lighter aromatic streams. Laboratory fixed-bed microreactor studies conducted with model feeds containing 0–15 wt% naphthalene have shown that the adsorption equilibrium constant for naphthalene hydrogenation is approximately one order of magnitude higher than that for methylnaphthalene dealkylation under the same conditions, although the exact ratio depends on catalyst sulfidation state and hydrogen sulfide partial pressure. These observations support the interpretation that excess naphthalene acts primarily as a competitive adsorbate and hydrogenation sink, rather than as an inert diluent in the feed.

Kinetic analysis of naphthalene-rich feeds requires a reaction network that distinguishes 2-methylnaphthalene dealkylation, 1-methylnaphthalene dealkylation, naphthalene hydrogenation to tetralin, and tetralin ring opening. The rate of 2-methylnaphthalene disappearance is often described by a first-order expression in methylnaphthalene partial pressure and hydrogen partial pressure raised to a power between 0.4 and 0.7, depending on catalyst sulfidation and temperature. The apparent activation energy for catalytic hydrodealkylation of methylnaphthalenes is reported in the range of 150–220 kJ/mol, while the activation energy for naphthalene hydrogenation to tetralin is lower, typically 90–140 kJ/mol. Because the hydrogenation activation energy is lower than the dealkylation activation energy, increasing reactor temperature improves dealkylation selectivity relative to hydrogenation, but the improvement is limited by coking. When the feed naphthalene content exceeds 8 wt%, the concentration term for naphthalene hydrogenation increases, shifting the product distribution toward tetralin at a given temperature. The selectivity response can be modelled by including competitive adsorption terms for naphthalene, methylnaphthalenes, tetralin, and hydrogen sulfide on the same active sites; however, published kinetic parameters for this specific configuration are limited and should be fitted to pilot-plant data using nonlinear regression rather than taken from light aromatic hydrodealkylation models.

When Feed Naphthalene Exceeds 8 Weight Percent in a Two-Stage HDA Unit

In a two-stage hydrodealkylation unit, the first reactor is operated at a lower temperature to hydrodesulfurise and partially hydrogenate the feed, while the second reactor is operated at a higher temperature to complete dealkylation and dehydrogenate tetralin back to naphthalene. Feed naphthalene contents above 8 wt% shift the first-stage hydrogenation equilibrium toward tetralin and decalin, increasing the second-stage heat requirement and reducing the overall selectivity to naphthalene if the second-stage outlet temperature is insufficient to dehydrogenate the additional tetralin produced. For this reason, two-stage units processing naphthalene-rich feeds often adjust second-stage reactor inlet temperatures to 650–720°C and reduce liquid hourly space velocity to 0.5–1.0 h⁻¹ to maintain acceptable tetralin conversion. The hydrogen quench flow between stages is also increased to limit the exotherm from naphthalene hydrogenation in the first bed; typical hydrogen quench rates are 10–25% of the total hydrogen circulation. Downstream separation of naphthalene from unconverted tetralin and methylnaphthalenes uses distillation or crystallisation; naphthalene recovery by crystallisation is governed by ASTM D1160 for boiling range distribution and by melting point measurements according to ASTM D1493 for solidified naphthalene. The presence of feed naphthalene above 8 wt% can alter the crystalliser feed composition such that the eutectic point with methylnaphthalenes is approached more rapidly, reducing naphthalene crystal yield unless the crystalliser temperature is lowered below 45°C. Published data for specific two-stage units operating with feed naphthalene contents above 8 wt% are limited, but process simulation studies using Peng–Robinson equations of state indicate that the reboiler duty in the naphthalene recovery column increases by 5–15% when the feed naphthalene content is raised from 6 wt% to 12 wt% at constant naphthalene recovery.

Selectivity Loss Mechanisms Cannot Be Separated from Hydrogen Partial Pressure Effects

The primary selectivity loss mechanisms associated with feed naphthalene contents above 8 wt% are competitive hydrogenation of naphthalene to tetralin, ring opening of tetralin to alkylbenzenes, and polycondensation to higher-molecular-weight aromatic coke precursors. Hydrogen partial pressure exerts a strong influence on the relative rates of these pathways; at hydrogen partial pressures above 35 bar, naphthalene hydrogenation to tetralin is favoured kinetically, while at hydrogen partial pressures below 20 bar, carbon deposition becomes more prominent. The optimum hydrogen partial pressure for preserving hydrodealkylation selectivity in naphthalene-rich service is therefore reported as a narrow band of 25–35 bar when the feed naphthalene content is between 8 wt% and 12 wt%, although the exact optimal point shifts with reactor temperature and catalyst age. Hydrogen sulfide is maintained in the recycle gas at 50–200 ppmv to keep the catalyst in a sulfided state and to suppress excessive hydrogenolysis, but excessive hydrogen sulfide concentrations above 300 ppmv can inhibit desulfurisation and dealkylation activity. The selectivity to naphthalene from 2-methylnaphthalene is typically 80–90 mol% in single-stage catalytic HDA when the feed naphthalene content is below 8 wt%; at feed naphthalene contents above 8 wt%, the same units often report a selectivity decline of 3–8 percentage points unless reactor temperature is raised or space velocity is reduced. Published data for feeds above 15 wt% naphthalene are scarce, and pilot-plant results indicate that the selectivity loss accelerates because the equilibrium concentration of tetralin at the reactor outlet rises sharply with increasing naphthalene partial pressure. In such cases, operators may reduce the feed naphthalene content by blending with a low-naphthalene aromatic stream or by lowering the feed endpoint so that the naphthalene concentration falls below the 8 wt% threshold. The trade-off between selectivity and feed cost is evaluated using calibrated process simulation models that incorporate reaction kinetics for methylnaphthalene dealkylation, naphthalene hydrogenation, tetralin dehydrogenation, and coke formation; these models are validated against plant data generated under ASTM D2425 and ASTM D2887 analytical protocols.

Operational boundaries in naphthalene-rich hydrodealkylation service include the requirement to pre-dry feed streams when relative humidity exceeds 60% because water accelerates catalyst support degradation and shifts the gas-phase equilibrium in high-temperature hydrogen service. The feed must also be kept free of oxygenated compounds and organic chlorides, which can form acidic species that corrode reactor effluent air coolers and promote catalyst metal sintering. In units where the feed naphthalene content exceeds 8 wt%, the combination of naphthalene with high-boiling unsaturated compounds such as indene and acenaphthylene is particularly unfavourable; these compounds copolymerise in the preheat exchanger and increase pressure drop, reducing hydrogen partial pressure and further lowering hydrodealkylation selectivity. Batch-to-batch variance in naphthalene content is controlled by online near-infrared spectroscopy or by gas chromatographic analysis of the feed tank, and the feed blending strategy is adjusted to maintain the naphthalene concentration below 8 wt% unless the unit is specifically designed for naphthalene-rich operation with continuous quench and shorter catalyst regeneration cycles. Materials selection for reactor vessels in this service is based on chromium-molybdenum steel grades such as 2.25Cr-1Mo or 1.25Cr-0.5Mo, with weld overlay or cladding where hydrogen partial pressure exceeds 30 bar; the Nelson curves in API 941 provide the basis for selection of pressure vessel steels in high-temperature hydrogen service. Transfer lines and letdown valves are heat traced to maintain metal temperatures above 85°C to prevent naphthalene crystallisation. Published data for naphthalene-rich feed effects on hydrodealkylation selectivity are limited, especially for feeds with naphthalene contents above 15 wt%, and site-specific catalyst deactivation rates must be established through pilot-plant testing rather than extrapolation from lighter aromatic hydrodealkylation units.