After switching from RLNG to Naphtha Feed, my Pre-Reformer DP began to rise. Is there any chance of leaching out of the HDS Bed Catalyst due to low sulfur in RLNG? Other than that, what are the possible causes of the Pre-Reformer DP Raising issue?

How to fix the fouling issue at the feed distributor on the kerosene side stripper? At the HCU fractionator, the kerosene was drawn from the main fractionator to the kerosene stripper before being run down to the product tank. We have had many problems with fouling at the kerosene stripper's feed distributor. We open the valve to draw more kerosene, but the kerosene rundown from the kerosene stripper does not increase. The rundown flow rate is decreasing over time. Should we install a strainer at the kero draw line before entering the kero stripper feed distributor? Or should we upgrade the feed distributor (now it's a pipe branch with an orifice hole)?

What is needed to scale up chemical recycling of waste-plastic pyrolysis oils to the 10-30% range?

How can unpredictable cycle performance be controlled when co-processing renewables?

Can you point out cases where hydrotreater catalyst regeneration met the refiner’s expectations?

What aspects of AI can be leveraged to optimise reactor and catalyst design?

Ammonium polysulphide (APS) is currently being injected upstream of the intercooler and aftercooler, along with wash water. In the original DCU design documents, the Licensor did not include this chemical dosing facility. However, when cyanide levels> than 20 ppm were observed in the interstage and recontacting drum sour water streams, APS injection was initiated to mitigate cyanide-induced corrosion in these sections. APS dosing was started to keep the cyanide concentration below 20 ppm. We noticed that some DCU units operate without APS dosing while still maintaining cyanide levels below 20 ppm. We would like to hear your expert opinion on the APS dosing. Could you please share the best practices followed globally in DCU units regarding the APS dosing?

The DCU Feed surge drum (508-V-001) is designed to mix the hot and cold VR streams and to provide emergency backup feedstock during feed failure scenarios. At present, we maintain a positive pressure of 1.5 kg/cm2g in this vessel by continuously supplying fuel gas as a backup. As part of process optimisation, we propose using nitrogen instead of fuel gas to maintain the same pressure. Kindly share your expert opinion on this.

Potential Carryover of Amine-Based Corrosion Inhibitor to Isomerisation Feed: Analytical and Operational Experience in a Light Naphtha Hydrotreating (LNHT) unit. A film-forming amine-based corrosion inhibitor is injected at the stripper overhead. The stripper bottoms are routed as feed to a downstream chlorided-alumina isomerisation unit. We are observing gradual catalyst deactivation in the isomerisation reactors. 

Known contaminants such as moisture, sulphur, oxygenates, and CO are within specified limits. However, we suspect that excess corrosion inhibitor dosing may be contributing to nitrogen carryover into the feed. Based on dosing estimates, nitrogen levels could be on the order of 1–1.5 wtppm, exceeding the feed specification limit of 0.1 wtppm. Total nitrogen analysis using ASTM D4629 has not indicated elevated nitrogen levels, which may reflect analytical limitations in detecting long-chain surface-active amines in light paraffinic matrices. 

We would appreciate insights from others who may have encountered similar situations: 
- Have you observed corrosion inhibitor carryover impacting isomerisation catalyst activity?

- What analytical methods have proven reliable for detecting or confirming amine-based inhibitor presence at trace levels in light naphtha? 

- Are there recommended sampling or extraction techniques to improve detection sensitivity? 

- What mitigation strategies (dosing optimisation, relocation, guard beds, etc.) have been effective in practice? 

Any shared experience or guidance would be greatly appreciated.

In our refinery, we have both DHDT and DHDS units, both utilising the same CoMo-based hydrotreating catalyst. The operating pressures are: DHDT at 68 ksc and DHDS at 37 ksc. In the event of hydrogen failure in the DHDT unit, how much can the system pressure be reduced, and what impact would this have on the catalyst?

I am operating an SMR unit with a pre-reformer (RLNG feed). The pre-reformer operating inlet temperature is 490°C. During plant start-up, when steam and feed are introduced, the pre-reformer temperature will drop. What are the consequences if the temperature falls below 400°C, and is there any impact on the catalyst?

We need support. We need to know the perfect value for the brine of the desalter?

Kindly clarify whether aeration of effluent water alone is effective in reducing ammonia concentration, and if so, to what achievable outlet ammonia levels (ppm) under normal operating conditions. Further, please share the recommended operating parameters for the aeration system, including, but not limited to, airflow rate, hydraulic retention time (HRT), dissolved oxygen (DO) levels, pH range, temperature, and expected ammonia removal efficiency.

Recently, facing issues in Desalter Operation: Desalter Efficiency (salt removal) is 95%, but Desalter Effluent (brine) colour is BLACK. Sometimes there is no oil carryover in the Effluent sample, yet the colour remains black. Demulsifier dosing rate has been increased from 5 ppm to 14 ppm. Mix valve DelP maintaining 0.9 ksc. Wash water injection was also maintained at the designed rate. Is it due to the presence of sludge in the crude tank, or is it due to another reason?