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Details Include any more information that will help us locate the issue and fix it faster for you. Thank you for submitting a report! Submitting a report will send us an email through our customer support system. Submit report Close. Recommended Articles Loading There are no references for this article. Why only water is getting carried over in huge quantities along with the treated gas?

Is the same observed in any high pressure amine absorbers. Why temperature profile in the amine absorber is in the opposite direction? Is this related to moving of heat of absorption profile from bottom to top? Further info: If there is a foaming in the column the following should have happened: 1. Fluctuations in absorber bottom level; not fluctuating 2. Severe fluctuations in DP across the absorber; not fluctuating 3. Improper stripping of H2S from gas; not observed 4. Amine foaming tendency test is also done and found satisfactory.

In the case of foaming in the absorber, amine should also get carried over along with water. But this is not happening.

As i said earlier amine content in the water sample is in the range of 0. What is the ideal delta temperature to be maintained between gas and amine? Direct stripping steam is used here. We used to get product diesel Cu strip result as 1b.

The following are attempted to normalize the section: a. Checked the steam line for condensation. No leaks c. Stripping steam increased gradually from 3 MT to 4. Withdrawn more distillate from reflux drum. I could not find any H2S slip in bottoms for both the conditions. Even i tried to simulate the column by taking out few trays.

There is no improvement. Please provide your valuable suggestions to improve it further. Also please provide reply for the queries given below: 1. Cu strip result definers with color.

But how much H2S will be there in product if the result is 1, 2, 3? How to find out whether there is flooding in the column. DP across the column is at 0. Will it vary severely if column is flooding? We have 2 reactor in series guard bed and main reactor. Guard bed reactor is loaded with demat catalyst. During start of run the delta t across guard bed reactor was 28 Deg-C.

But during six months we are observing that delta t of guard bed reactor is gradually reducing and had reached 6 Deg-C. Guard bed reactor delta p is normal and is around 1.

Also we observed that all guard bed rection had been shifted to main reactors and we are getting high delta taround 25 deg-C in 2nd and 3rd bed of reactor.

What is causing this? We experienced a tube rupture in our vacuum column re-boiler due to over heating and once we introduced emergency coil steam in re-boiler tubes, a major fire broke out in the furnace, leading to complete destruction of furnace.

Kindly comment on this as it is not clear to us whether it was a right decision or should we wait to cool down the furnace? Then superheater is used for heating isomerate outlet temp. Is process disturbed and will it poison the catalyst? Why can't we use superheater initially from 50 degc to degc? Is there any problem of heating liquid isomerate or it can lead to coke? Is superheater used for only vapor phase not liquid phase?

I want to install Hijector for boosting the pressure. Can anybody suggest what will be the pressure of my motive fluid, whether such scheme will work and, if not, what are the alternatives? The filter element are wedge wire with 75 microns. We are doing some studies to identify the origin of the problem.

VGO has ppm of solids, which are mayoritary inorganics particles. The higher temperature the higher unstable is the mix, and the asphaltene precipitate at lower HCGO percentage. My first question is if somebody has experience of this sort of event? We think the solution is not to increase the filter area, but eliminate the problem at its source, to reach a HCGO cleaner wiht less asphaltene content. My second question is related to the effect the asphaltene precipitation with the temperature.

I thought that the higher temperature the lower precipitation but we have seen the oppsoite effect. What happens when number of trays is increased? There is a project under consideration where hydro-treated jet will be produced through a Hydro-desulphurization Unit. The hydro-treated jet with have anti-oxidant injected into the rundown to storage and will co-mingle with straight run Merox treated jet before entering the same storage tank.

To achieve this tank mixers or a re-circulation system is being considered. Based on API , section 4. Other important considerations are the co-mingled jet will be stored in fixed roof style tanks. Secondly there is no anti-static additive injected in the tank at this point as it gets added at the truck rack and marine terminal.

My questions are: 1 Do other refineries typically use mixers or a recirculation system on jet tanks when co-mingling different types of jet? With all product tanks on-site the low pump out level is always above the mixer elevation and a low alarm ring in when approaching low pump out.

But we are now facing problem of high column pressure of around 1. For that column remains upset most of the time as naphtha is not removed properly. Our product flash point is also found lower than design. For that our high naphtha production always remains a concern as light end section found upset.

It also observed that our column feed inlet temeprature always kept slightly lower 0C than design of 0c. Presently we bring down the column pressure to nearly design pressure of 1. Now I just want to know can we now keep column top temeprature 93 0c as per design for less naptha production and also meet the flash point requirment of products by increasing the column feed inlet to design 0c.

On what design basis would a vendor recommend pilot operated safety valves in a refinery? How are they different from normal spring type PSV? From what I have read obviously not enough , my logic dictates: 1 Unsaturated hydrocarbon molecules are more volatile than that saturated ones, thus will have higher RVP 2 Lighter hydrocarbon molecules are more volatile than longer chains, thus will have a higher RVP. In other words, which has higher RVP - an unsaturated aromatic benzene molecule or a saturated paraffin pentane?

I am actually in finance trying to figure out product properties. The questions driving me nuts: 1 Why does alkylate which has more saturated content have higher RVP than reformate? Last time during pigging activities we receive many sludge from offshore, and now all this sludge is settled down inside Rich MEG tanks.

Any ideas how to improve situation with Rich MEG tanks? Maybe clean Rich MEG using hydro-cyclones, or any other equipment? Any links to useful equipment to be installed, or to similar problem anywhere?

TQVM in advance. This make-up gas can be used as hydrogen for the DHT to keep it running for the duration of the whole hydrogen plant outage. What is the impact on -Reaction section -Product quality -Recycle compressor?

I am plumbed into the unit with my diesel powered pump that has taken place of two electric drive pumps that have failed for undisclosed reasons to me at this time. If anyone has any experiences with these pumps could you enlighten me to the hazards involved, the use in process, and any down stream side effects on a refinery when they are out of service?

Also I was told that within twenty minutes of shut down on their pumps that their unit would cease to function due to salt build up.

Please recommend some other alternative for the removal of the marcaptan sulfurs so that it can be directly routed to the CRU feed pool. How do I calculate the dew point of overhead vapors? In one of our refineries we have detected high sulphur content in LPG from crude distillation. The scheme is as follows: Crude distillation - Gas concentration unit - Debutanizer - Amine absorber - Merox extractive. We have also seen some unexpected behaviour with these species: DMDS increase through the amine plant DMDS higher in the outlet than in the inlet and decrease in the Merox unit.

The same with DMS. But sometimes we have also observe that DMDS decrease in the amine plant?? DMDS could be re-entry sulphur in Merox, but we have observe it in the inlet of amines and Merox It seems that both compounds come with the crude - Could be DMDS come from oxidation of methylmercaptan in the topping, Gascon or amines where there is not Merox catalyst if oxygen is present in the LPG? I do not think so. Are these compunds partially soluble in NaOH or could be removed in the sand filter?

What is the normal or recomended concentration of disulphides in regenerated NaOH? Improves copper strip corrosion test adding corrosion inhibitor My questions are: - Could the low level of mercaptans present cause a failure in copper corrosion strip?

We are having 24 filter for removing contaminants including one backwash filter. Our filter dp during steady state operation is 0. We often face problem of high pressure drop of 3. I am not able to find out the root cause for high dp across feed filter during such activity in crude.

So can you tell us the what is the difference between Stripper and Splitter? Additional info:- Stripper and Splitter both are having reboiler for temperature controller. Same we checked in product fractionator heater. But after 2 to 3 days it was observed that product fractionator heater pass 1 convection to radiation vibration is significant.

We reduced the throughput to KBPSd. But still vibration persists. Can anybody guide me about sudden increase in vibration? The compressor was tripped off due to the false ESD activation of high level switch on the suction K. O drum. Recycle compressor started up and started increasing the reactor inlet temperature.

During the DHT unit start up, higher hydrogen consumption of 1. My question is where is go the higher hydrogen consumption?

But the sulfur content of C4 is not decreased. Investigating the cause of amine absorber foaming, we find the significant change of amine absorber condition. There is amine carry over to overhead LPG side in amine absorber. Second is LPG carry under to rich amine side in amine absorber. Rich amine goes with LPG to amine flash drum before amine regenerator. So the pressure of amine flash drum sometimes rise to almost drum design pressure.

Finally we replace the activated carbon filer in rich amine side, but there is nothing wrong in amine quality. After that, Inlet and outlet amine flow is same and the delta P of amine absorber increase to normal condition We wonder why LPG absorber goes back to the normal condition after replacement of rich amine filter.

Could you explain the reason for this phenomenon? If amine quality is main cause, could you recommend the new guide of amine or other countermeasure? Hydrogen sulfide is removed by counter current of amine solution and the LPG leaves the top of the column and flows into the amine settler D and rich amine is leaves the bottom of the absorber to amine regenerator.

LPG flows into the caustic prewash drum D for removal the last traces of H2S not removed in the amine absorber. D is Caustic Settler. One 35bar is producing Euro III diesel with a sulfur specification of ppm maximum and other bar producing Euro IV diesel with a specification of 50 ppm sulfur.

Now I would like to know what type of catalyst to be used. Co-Mo catalyst or Ni-Mo? Moreover kindly explain the basis for choosing the Catalyst-type. My question is if recycle gas compressor trips and our feed pumps remain running due to faulty logic, what will happen to reactor?

And can we run the feed pump for cooling of reactor without recycle gas? Root cause? Splitter 52 trays Stripper 25 sieves trays. And if I would like to revamp this unit for a product with 0. We have 1 reactor R1 with 1 bed of catalyst 18m3 catalyst in 27m3 reactor. I think we should install one more reactor.

May you have any advice for our revamp? Additional info: Of course that Case 1 is traditional process revamp. But I have just read an article from Chevron, about their process revamp as Case 2. I think in case 2, R2 is existent reactor and R1 is new one because R1's volume needs to be bigger than R2 This article named "Hydroprocessing upgrades to meet changing fuels requirement", Jay Parekh and Harjeet Virdi. Is it O. Field observation for any abnormal sound across the NRV was checked and found normal.

Actions Taken: FT installed at the Quench valve was also drained and purged and found no error. Unit load reduced to turn down ratio. Your expert opinion and guidance is requested on the Issue. Any thoughts on reasons and solutions? Since start of Unit the filter choke again and again. Some times Unit thruput is reduced as these filter elements have to be manually cleaned and the cleaning interval reduces to less than a day.

Our crude composition changes with change in crude tank and sometimes residue is there in Gas Oil. The filter elements are chemically cleaned and the interval increased but problem remained. What are other refiner's experience? What may be other causes of Filter elements chocking? Can high amount of residual water can choke filter elements? DP increased to mmh2o from normal one mmh2o. My question is, is high paraffin in feed is the problem or some another causes, if yes, then how?

Additional info: Tank feed bromine was analized and we got 1. Kindly suggest what may the probable reason for such a trip without any prior alarm? Compressor is steam driven.. Is the trip incident caused by variations in gas molecular weight?

Is it possible that surge conditions occur due to comparative lighter gas handling than design operating which ultimately force the RGC to trip? My feed conditions are : Temp at battery limit: Deg C, pressure: 6. Density: 0. Kerosene is being filtered by two basket type filter having mesh one stand by Filter temp around: Deg C. We are facing a problem of frequent filter chocking, but filter element is clear, no dirt, no scale, no corrosion particles, you can say crystal clear like clean filter element, still having high DP.

Additional info: Filter is getting chocked frequently. Once filter got chocked 16 times in 2. Dirty filter baskets are being cleaned by hydrojetting and followed by steaming. Original design was of 25 micron mesh , but because of frequent chocking filter mesh has been changed to 74 micron mesh temporarily. Filter element is of stainless steel. Till date no adverse effect observed in reactor DP or heat exchanger fouling. LK feed is straight run from crude unit, no feed from tankage.

These are recovered in two columns under vacuum. Unrecovered stuff is sent to Utilities as liquid fuel. Anti-oxidant injection is done in the Ethylene unit as Pygas contains precursors such as dienes which can lead to polymerisation.

Recovery unit was operating steady, without any problems, for 8 months. Now for some reason the frequency of choking of the strainer of bottoms pump of the last column has increased dramatically. Also, we are experiencing frequent choking of burner guns. Material found is coffee coloured granules which become powder when subjected to pressure.

Trying to understand root cause. Not much has changed in terms of operating conditions. Very few component analyses are done in the whole system and not much information is available. Hope to get some inputs based on experience in similar units. We are continuesly loosing Reactot Delta T.

At SOR it was deg C. What are the possible causes? However, stablizer overhead gases has also increased extensively. Some opinions arises that there may be the leakage in Combined feed exchangers of Platforming section. But, we are unable to detect this leakage during plant operation. Please mention, how we can detect this leakage during plant operation and secondly, please also describe that what may be the other reasons of such decreasing trend of delta T i.

Additional info: Its again me who put up the questions. Can some one suggest if it is possible and what are the critical parameters for it? We need to estimate the cost of the unit and its facilities like the vacuum tower and the vis-breaker. How would you suggest we get a rough initial estimate of the costs involved? The first tower operates at 7 KSCg pressure and second tower operates at 0. Recently we have encountered a strange problem. The color of the stripped water is milky white and also looks hazy.

The overhead temperature of the second tower is running high, C Normal is 90C. Please suggest some solution. In the startup procedure, we were told that feed should be cut in at a reactor bed temperature of C. But our current catalyst supplier has suggested that you can cut in feed even at - C. I just want to know what will be process implications if we cut in feed at C or C or C. Our diesel feed API is 33, feed sulfur is 1.

One more thing: what will be effect if we run the high pressure separator at lower and higher temperatures. It happens some times because of Fin-Fan cooler problems and climatic conditions. However measuring the VOC content at breath out shows as high at ppm. The internal roof rim seal was replaced and produced only minor improvement. Is there any plant try to install vapor recovery unit to reduce these emissions?

Is there any regulation which requires the benzene tank to be equipment with close system? In Sour water strippers why is chimney tray provided? I would like to know how much should be the wash water flow.

I also look after Sour water stripping unit. Why are the feed valves to SWS stripper located near the tower. Is there any specific reason for it? The same is the case with Amine recovery unit stripper. But, in some equipment, we use turbine pump as primary pump and backed up by turbine pump as a spare pump. This pump transfer the bottom of low pressure separator liquid hydrocarbon to debutanizer. I also found a pump configuration where both the primary and spare pump are turbine pump.

This pump is diesel pump around hot wash. Do you know what is the reason behind these configurations? And how is it processed? Is there any possibility of sulfur stripped out from liquid phase sufided catalyst? Pl let me know whether dew point of stripping steam used in distillation column depends upon partial pressure of steam in the total vapour mass flow going out in the distillation overhead. Pl let me know if you need more data to answer my question. Our distillation column top operates at 1.

In some hydrogen plant it is mixed only in reformer inlet. What is the advantage of that? Currently due to some problem in the electrical heater in the Fuel Oil Circuit we are using only fuel gas.

Some days back inspection department reported a much higher skin temperature in the radiation section of the Furnace. The same report was also upheld during various cross-checks by other departments. Could this be due to the reason as we are not using Fuel Oil? If so, then could somebody explain? Are there any other models available in the market? Are there any tools which can be helpful in daily monitoring of the hydrotreating reactors? The product naphtha fails due to colouring.

But the other products are passing all the required test and parameters. Has anyone faced such problem; if so what could be the reason for naphtha product alone getting coloured? Recently our unit tripped because of some strange problem. I request all to suggest a reason for the problem explained below.

We have one centrifugal Recycle gas compressor RGC and two reciprocating make up gas compressors MGC one running and the other stand-by. As per the regular change over of MGC we tried to take the other one in line and spare the running one. As the discharge pressure of RGC reduced the discharge flow also reduced.

I didn't understand why the discharge pressure of RGC came down. The suction flow was Tons per day and when the RGC discharge pressure dropped, the suction flow also dropped to Tons per day.

I have noticed that whenever the mole ratio is increased by virtue of increase in recycle gas flow, sum of delta Ts across reactors drop, especially across Rx 1. It was also observed that just prior to surging the total flow at the inlet of the RGC was also increasing. We have got an amine column at the inlet of RGC suction after HP separator to reduce sulphur loading.

But now due to some constraints the amine flow had to be reduced. Can anybody explain the phenomenon? Please recommend some other alternative for the removal of the marcaptan sulfurs so that it can be directly routed to the CRU feed pool. But you have to look for a good adsorbent in the guard as life of the adsorbent is also an issue. The location of the guard should be made based on a number of issues.

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