Q & A > Hydroprocessing
Date  Replies
19/05/2017 Q: In VGO hydrotreater unit, cracked feed processed from coker unit which is mainly HCGO and HHCGO (Heavy Heavy coker gas oil). If HHCGO end point increased from 570 deg C to 600 degC, then how severe its impact on VGOHT operation and its catalyst performance?  
16/05/2017 Q: in our C5/C6 Isomerization unit, the reactor effluent after feed/effluent heat exchanger has 43% liquid at 1040C. this effluent is being cooled to 400C using air finfan cooler and water trim cooler. Hydrogen rich gas is separated from liquid hydrocarbon in a separator at 18 kg/cm2 and used as recycle gas. We normally maintain H2/HC = 1.5 mole ratio.
We want to install a hot separator at the outlet of feed effluent exchanger to separate the condensed hydrocarbon and the remaining gases and uncondensed vapour will be further cooled in air fin fan cooler and water trim cooler. Liquid hydrocarbon from hot separator and cold separator will be mixed and sent to stabilizer to remove C1-C4 components.
1) I want to know whether this type of hot separator is feasible and did any refinery has this type of system.
2) How much will it affect recycle gas purity.
14/05/2017 Q: We are using AXENS Catalyst RG-682 (Naphtha Reforming) and HR-538 (Naphtha Hydrotreating).
Now, my question is what is the actual life time of these two Catalyst? How many times it can possible to regenerate RG-682 & HR-538?
12/05/2017 Q: I am process engineer of Naphta Hydrotreating units in a Refinery with capacity of 350000 bbls/day. In this moment the refinery it has in storage tanks gasoline with higher weight percent in MTBE above 2 % weight. However this gasoline will be send to co-processing in VGO Hydrotreating unit (FCC Pretreatment) where according with our evaluation it is possible co-processing this naphta by dilution in the reactors taking in account because HDO reactions, will have more heat release, hydrogen consumption and water generating.
So we could hydrotreated this naphta in VGO hydrotreating unit, however the quantity of oxygen compounds like MTBE was about 20-40 ppm. This naphta is send to the naphta fractionator and will be separate in Light naphta (LVN) and heavy naphta (HVN), and the HVN stream is send to NHT unit. But the problem is, according with opinion of process engineer of NHT unit, the feedstock to NHT should not have oxygen compounds (0 ppm) in order to avoid fouling in the preheat train by gum formation, but the catalytic scheme installed in the reactor of NHT there are CoMo and according with my evaluation, after dilution of this naphta previous introduce to the NHT, the concentration of ppm in MTBE will be lower and i have read and studied, that in NHT units, the most of oxygen compounds are converted around 90 %, so in my oppinion would have not problems in send this feedstock to NHT unit.
My question is, does it exist a maximum concentration in ppm of oxygen compounds like MTBE in feedstock to naphta hydrotreating units? and which will be the main impacts in point of view catalytic and heat transer in the preheat train?
10/05/2017 Q: We have platforming heaters tubes of alloy 2.25 Cr-1Mo (vertical tubes) .
The maximum design skin point temperature 595c while the limiting design metal temperature 650c as per API 530.
What is the maximum temperature we can reach it above the maximum design and below the limiting design to avoid the oxidized of the metal?
02/05/2017 Q: How can I calculate Reforming Heat of reaction and reactor Delta T. From the catalytic reaction guideline I know that the Napthene dehydrogenetion heat of reaction -50 Kcal/mole. Now I want to calculate reactor delta T.
Additionally I know the reformer feed flow rate, feed detail hydrocarbon analysis, feed density, feed molecular weight
In practical operation, we have three reactor in series, 1st reactor delta T 117F, 2nd reactor delta T 48F and 3rd reactor delta T 16 F; Now I want to calculate this delta T in theoretically. How can I prove/calculate that this practical delta T as like theoretical?
19/04/2017 Q: I currently working with naphtha hydrotreater plant we have splitter column which split only light and heavy naphtha.in our reboiler (fired heater) of that column convection inlet temperature is 155C and convection outlet is 181C further it goes in radiation inlet where temperature same as a connection outlet but outlet of radiation is 161C. I don't understand that why temperature in radiation zone decrease compare to convection outlet.
30/03/2017 Q: What is Wild Naphtha and why is it so called? (5)
23/03/2017 Q: In our hydrocraker unit in first stage we have VGO and HCGO. We maintain ratio 70:30. If we increase the percentage of HCGO there is any effect in reaction and hydrogen consumption? And how to maintained? (6)
20/03/2017 Q: We are using Compressor Type WS2/180-A1 for our Naphtha HydroTreating recycle gas compressor. During Wash water injection at the up stream of NHT reactor effluent separator, water is carrying over to the suction of Compressor. As a result of that compressor discharge, flow becomes low and load current also becomes low which suggests one of the load valve is not functioning due to dirt in water may have choked the load valve. After cleaning the load valve, flow and current becomes normal again. This problem is being faced recently when some salt has formed at the downstream of NHT reactor effluent which is dissolving with the wash water and carrying over to the suction of compressor. In addition, we are facing this problem after 3.5 years of plant life cycle. My question is whether it is happening only for dirt in water or load valve's diaphragm and O ring get old or spring has lost its tensile strength. Whether the Compressor type WS2/180-A1 is designed to handle some liquid or water as it is designed for recycle gas of NHT unit? FYI- 1. Wash water injection has been continuing in a regular interval of 15 days since the inception of plant start-up 2. Salt formation at the suction strainer is experiencing currently.

20/03/2017 Q: Currently Chloride content in our straight run Heavy Naphtha (HN) feed is 800 PPM. So, please tell me the maximum limit of Chloride content in HN is acceptable for NHT unit to minimize the chloride corrosion and salt formation in the system (for Naphtha Hydrotreating Unit). (5)
20/03/2017 Q: We maintain 180-200 ppm H2S gas (measured by Drager Tube CH29101) at the outlet of NHT recycle gas compressor discharge to keep the HR 506 (Co Mo) catalyst in its active form or sulphided form.FYI, our HN feed contain very low sulfur which is Doctor test negative. For this purpose,we are adding 2.8-3 kg pure DMDS with the feed which can be able to produce more or less 0.095 kg H2S in the recycle Gas but from the design data book calculated value is 21.99 Kg H2S (provided recycle gas flow 467 kg/hr) which is 47,087 ppm.So, how the calculation has done and how it becomes 180-200 ppm H2S in the recycle gas. Please show me the basis of calculation. (4)
20/03/2017 Q: I am in charge of DHDT (produce 10ppm ULSD) catalyst performance monitoring, recently WABT was increased fast out of my expectation, meanwhile ULSD color was getting worse (DHDT has decolor reactor), observed cut point (T95) of feedstock SRLGO had a fluctuation in last few weeks, what is the possible cause of this abnormal WABT jump? (5)
15/03/2017 Q: What is LSHV. What is effect if decrease or increase in the reaction .And how to maintain it? (3)
15/03/2017 Q: In our NHT unit, tube material of stripper Column Overhead air cooler is SA-179 which is low carbon steel. So, if we use type SS 321 instead of SA-179 then will it be more sustainable in the wet H2S and wet HCl environment? (3)
14/03/2017 Q: Is there any nitrogen species that may be present in LVN but is not present in HVN? We are detecting high nitrogen content in LVN but not in HVN. Also, our sulfur content is low. Even though we may see nitrogen in HVN (poison to reactor), the endotherm of the reforming reactor is not affected. Are there nitrogen species that can be detected by NSX but is not readily available for breaking down/reaction? (3)
11/03/2017 Q: Please tell me the normal PH limit supposed to have at the NHT Stripper Column Overhead drum water boot sample. Currently, we are getting ±1 which is very low. Please enlighten me whether it could be the effect of wet H2S alone or it could be the cause of combined effect of wet HCl and H2S which might carry over during the wash water injection from reactor effluent Separator drum to the Stripping Column Overhead Drum.
FYI- 10 wt ppm CHIMEC 1044 is being injected at the inlet of stripper overhead air cooler continuously.
01/03/2017 Q: We have tried several times to continue inject wash water @ 80-100l/h at the upstream of air cooler of NHT separator drum but every time recycle gas flow becomes low remarkably due to suction valve get choked which seemed some entrained water and salt carried over and blocked the suction load valve of Compressor. Therefore, please advise us what could go wrong and how to solve the problem? Moreover, we have increased the temperature after Separator aircooler (A-201) to avoid the tube corrosion which reduced the DP fluctuation across the reactor R-201 but frequency of salt formation at the suction strainer was increased and it seemed de-sublimation point of salt attained at the downstream of A-201 which is close to the suction of recycle gas compressor. However, last couple of days we did not inject the wash water and monitored the rate of salt formation at suction strainer and found very insignificant amount of salt that could be the reason the de-sublimation point shift and not coming close to the suction of compressor for the time being. But, Today again get problem just after injecting only 50 litre wash water but no salt is found at strainer which seemed water carry over and choke suction valve again. For your info, we have drained water from the separator boot continuously- is there any problem with the water injection point? So, should we shut down the plant and need to do steam cleaning whole NHT Circuit to eliminate all the salt from the system? (4)
11/02/2017 Q: How many Vacuum Residue Hydrocrackings are in operation worldwide?
What is the maximum conversion achievable (ie minimum bottom product after the atmospheric and vacuum distillation of the reactor effluents) in these plants?
11/02/2017 Q: In hydrocraker unit in make up gas compressor I absorb first stage KOD LT and LG connected to ATM but second stage KOD LT and LG connected to flare. Is there any reason for that? (1)
07/02/2017 Q: We are treating Heavy Naphtha to produce DSN in Naphtha Hydro-treating Unit. The effluent stream from reactor bed flows to a Separator Drum after cooling through an air-cooler. As per the recommendation at the up stream of air cooler, we are giving 350 litres/ hr wash water after each 15 days interval. Even though, instead of NH4Cl salt we are getting Iron Chloride salt at the suction of Compressor strainer. Therefore, frequency of changeover of compressor has increased remarkably. In this case, what should be the preventive measure for the air cooler as corrosion product is generating. This problem has been observed after 3 years of operation life. Can you share your experience please? (7)
04/02/2017 Q: We have a semi regen platformer unit and catalyst type UOP-R56. During the regenerating after the oxidation step we stopped at this step The problem of providing electricity occurred and we did not complete the rest of the steps, now the reactors under positive pressure by hydrogen reformer purity of H2 70%
Is this situation adversely affects the catalyst after the oxidation step?
And what needs to be done in this case?
04/02/2017 Q: We use straight run Heavy Naphtha from Tank which has floating roof. Now, my question is what should be the limit of moisture or water content in wt ppm level before entering Naphtha Hydro Treating Catalyst bed. Currently, we have 104 wt ppm moisture in HN feed can this quantity could be the cause to increase the pressure at the inlet of Reactor suddenly to create a peak of DP across the reactor bed? Currently, DP across the catalyst bed is 3 and it has reached gradually after the 3 years of operation cycle but become confuse about the sudden peak. So, please help. (5)
26/01/2017 Q: Current DP across the Naphtha Hydro-treating Reactor is 3 and it is giving peak frequently without giving a gradual increase of DP. Why is the pressure at the inlet of Reactor is suddenly increasing for the few second but coming down to previous position again? However, we are thinking Catalyst top bed is blocked by coke, iron scale and other metal contaminant. Besides, we are getting some Ash color and salt type dirt at the fine mesh filter at the suction of NHT compressor. So, i need to know what could be the probable composition of the dirt? For this, we assumed that it could be fouling product from the separator drum or dust of alumina ball from the bottom of Catalyst bed. Now we are facing frequent compressor changeover and cleaning of suction strainer. So, please share with me if anybody have similar kind of experience and help me to sort out what could be the composition of dirt. (3)
24/01/2017 Q: How can we reduce the excess hydrogen content in a hydrotreater off gas stream (off gas from reactor effluent separator/flash drum, stripper off gas)? (3)
24/01/2017 Q: What is the difference between hydrogen consumption ratio (makeup gas flow rate/ oil flow rate) and recycle gas-to-oil ratio (recycle gas flow rate / oil flow rate)? What will be the effect on the amount of make-up hydrogen required if we reduce the recycle gas flow rate? (5)
12/01/2017 Q: We have an issue with our RG compressor in hydroprocessing unit. The flow of the RG compressor (Reciprocating single stage) decreased all of a sudden while taking scheduled changeover of the RG compressor. Flow went from 22000 Nm3/hr to 17000 Nm3/hr. We thought it may be due to some problem with the compressor that we took changeover so after a while again C/O was taken and the previously running compressor was taken in line but the result was same.
The other process parameters are same like the HP separator pressure and temp remains same. the power consumption of the compressor got decreased from 125 KWH to ~85 KWH.
We have checked the spillback valve of the compressors and cleaned the strainer of the compressor but no results obtained.
MUG compressor flow remains same and the MUG discharge is in the Fin Fan outlet to HP separator. The purity of the RG remains same. And one thing to be noted is that the D/C pressure didn't change, the S/c pressure got inc by 1-1.5 KSC at the time of C/o but now its normal.
The RG suction drum is common for the both compressors. It was noted that for past few weeks the suction drum level got inc and draining was done a bit frequently, so could carry over of the Hcbn liq to the compressor cause this reduction in flow?. The one imp thing to be noted is that both the compressors are not developing the flow. the flow has got dec to 17000 nm3/hr and is constant at this value, no further reduction is there, it has been a week since the event.
(suction valve plate should be OK as the s/c temp is not inc) Pls give your valuable suggestions.

Thank you Mr. Rupesh for your valuable time and comment. Actually we have checked the unloaded valves after taking the compressor for a maintenance and checked it at site (operational wise) but the working was found normal. Actually after doing a thorough study we have come to an conclusion (as of now) we found out that there is a increased liquid carry over to the compressor. We have checked the strainer in short durations. And the possible cause could be some internal integrity damage at RG suction drum or High pressure separator (suction point of comp). So this activity of integrity check could be done only in available shutdown activity. As the gas to oil ratio is less than earlier but not low at alarming level, management took the decision of waiting till planned shutdown. P.s: I will update you once the activity is being done.
12/01/2017 Q: We are currently designing a new grassroots unit for diesel hydrotreater (DHT). There are 2 different opinion when it come to hydrogen mixing point: either it is mixed before or after combined feed exchanger (CFE) .
The view for the mixing point to be after CFE have concern about polymerization or faster fouling inside the CFE while the view with mixing point before CFE saying the impact will be totally the opposite.
What is the basis/philosophy for DHT design on where to put the mix point?
31/12/2016 Q: In our DHT plant we are preocessing Light gas oil and Heavy gas oil. Diesel product 95% recovery maintain at @355degc. Diesel product density maintain @822. We are optimising kerojet production over diesel in condensate fractionation point. As we draw more kerojet at condensate fractionation point Diesel become little heavy. After reactor outlet in stripper we are facing problem of less reflux and resulting in higher overhead temperature. Our diesel product total sulphur content is maintain in range of 4-6 ppm. How to increase reflux flow in stripper with kerojet maximisation at condensate fractionation unit? Is there any relation between total sulphur of diesel product and stripper reflux flow? (4)
27/12/2016 Q: We have found black solid deposits upon cleaning of our CCRU Net Gas Compressor First Stage Strainer. Upon analysis of composition, we have found that the sample contains hydrocarbon plus a significant amount of Chloride and Iron and with traces of Aluminium, Magnesium, Silicon, Phosphorous and Sulfur. What could be the source of these black solid deposits? (6)
23/12/2016 Q: Why do we need Forced Draft and what is its function? (6)
25/11/2016 Q: We have a semi regen platformer unit. After regenerating the catalyst, unit was started up and observed that reactor 1 delta temperature (we have 3 reactors ) is low and almost same as reactor 2 delta temperature. And the Platformate RON is also low at about 92 whereas the expected RON at start of run is about 94. This is the first time we experience this kind of a behaviour. There was no sulfur ingress or any other change in the feed stock (Heavy naphtha). The issue seems to be in reactor 1 and other two reactor delta temperatures are similar to the previous cycles. Can anyone help me on possible causes for this issue?

Additional info:
Appreciate the valuable comments. To add some more information, We do process only the virgin naphtha. Sulfur and nitrogen are less than 0.5 ppm as required. Feed is from Murban crude which does not have high metal content.However, it appears that the reactor 1 delta temp is in decreasing trend and reactor 2 delta temp has increased slightly. H2 purity also low and gases and lpg are on higher side. PONA test indicated that paraffin conversion is low compared to previous cycle SOR conditions. Catalyst samples were checked after oxidation and reduction steps. Appearance of reactor 1 samples were better compared to reactor 3 except the reddish rust coating in reactor 1 was high. But, this coating in reactor 1 was same as in previous cycles samples as well. Also, we are maintain little higher cl level as this is the 9th cycle of the catalyst and expect lower cl retention in catalyst.

New info:
By today, the reactor 1 delta temp has further decreased and reactor 2 also is seems to be in decreasing trend. A new information is that there was a upset in recycle gas compressor during regeneration (oxidation step). Seal oil loss increase due to seal damage and both the seals were replaced along with damaged bearings. But, labyrinths haven't changed. And we observed some seal oil in compressor discharge line ( some drops of seal oil leaking at the compressor discharge isolation valve flange). Therefore, we checked at casing drain line and another low point in discharge line. First day (two days ago), we found some significant amount of oil collected. After draining them, the next two days found only small collection. Anyhow, now we are doubtful about a lub oil contamination. In addition, can there be a possibility of channeling the first reactor? Because if so we have to dump the catalyst during regeneration. Another thing is that during regeneration, there was a strong detection of hcl in reactor 1 outlet ( more than 500 ppm).
24/10/2016 Q: Has there been a case in any Refineries where the backwash stream (post backwash) is routed back to the cold feed tank instead of slops. The idea is to minimize sloping and reprocess it. (4)
19/10/2016 Q: What is effect of Pressure, temperature, LHSV and H2/Feed ratio on gas oil sulfur removal? (4)
13/09/2016 Q: What steam stripping is preferred method for Diesel streams and re-boiler type of stripping is preferred for Naphtha? (4)
01/09/2016 Q: Can fresh hydrogen be used in Hydro-cracker reactors as a quench media in case of recycle hydrogen failure? (5)
31/08/2016 Q: We are facing color problem in our Crude Distillation Column, during processing Color of Naphtha noted +17 instead of +30 what are the possible causes of this color problem.
Also color of Kerosene become off spec +10 instead of +30 or +28 what are the possible causes of this problem.
We checked all the heat exchangers for Possible leakage but no leakage observed.
22/07/2016 Q: What factors affect aromatic saturation in hydroprocessing of coker heavy gas oil? How can the aromatic saturation be minimized for heat balance benefits in the downstream FCC unit? (6)
19/07/2016 Q: Why is a rundown block-in timer required in modern hydrogen compressors for hydrotreating applications?  
15/07/2016 Q: IN our HCU Recycle Gas Compressor Turbine, make BHEL from one month Turbine Front & Rear journal bearing vibration suddenly increase from 15 micron to 80 micron for about 10 minutes and came back to normal during this period turbine bearing temperature also increase 3-9 degree centigrade. Please suggest the root cause. We do centrifuge of lube-oil once in 24 hors about 2-3 hours and lube oil also replace in december 2015. (1)
06/07/2016 Q: I am working in hydrocracker unit. Since two month LPG is getting failed due to positive H2S and in copper corrosion test. We are maintaining debutanizer bottom and top temperature 186 & 78 degree C and pressure 16.5 kg/cm2g. Our LPG r/d flow is10-12 m3/hr while lean amine flow is 20-25 m3/hr. When we check caustic strength found 20-22% which is quite normal. Generally we change caustic in 5-7 days.
Our system is like this LPG comes at top of debutanizer and goes in amine absorber at bottom and lean amine mix at top of absorber after amine wash goes to water wash amine settler where water circulation is done by pump before going to water wash LPG and water is mixed and mixed in mixer. After water wash LPG goes to caustic wash tank through nozzle but no circulation pump is there (Tank). After caustic wash LPG goes to sand filter and then goes to Deethaniser for FG removal and cooled in Water cooler and goes to LPG storage tank. Please suggest the solution.

06/07/2016 Q: At present in our VGO hydrotreater we have conversion of around 19.5 % to liquid distillates. we are facing problems with low recycle gas purity: 84-85% C1 and C2 content is high around 13%. What could be the possible reasons? We have recycle gas scrubber after the cold separator. (4)
06/07/2016 Q: We are working in a hydrocracking unit and since we start up, we haven't had a good copper strip corrosion data in the light nafta, we have 4a and the best we have achieved consistently is 2b. We lowered the pressure in the main stripper from 125 to 115 psig, increased overhead temperature from 282F to 292F and increased the stripping steam from 8000 lb/h (design) to 9600 lb/h. Also in the debutanizer we have drop the pressure from 160 to 140 ans still we don't have good results. What can we do in order to reach the nafta copper strip corrosion in 1A?


Thank you for your answers, I checked the steam and I saw it is 175 psi and 390F, so we are going to heat up more the steam and we are going to try increasing more the flow, but what could happen if I increase it too much, maybe the control valve 100% open and still not get the copper strip corrosion ok?
27/06/2016 Q: In case of heavy residue upgrading, we are encountered with vacuum residue as feed. The main features of this feed especially about contaminants and problematic materials are as below:
Total sulfur>4.5 wt%
Conradson Carbon >25 wt%
Ni+V >500 ppmwt
Nitrogen ~ 1 wt%
We have two cases for VR upgrading project, One is RCD+RFCC and another is HOIL(Hydrocracking)+FCC. Both of these cases use huge amount of fresh catalysts because of high possibility of catalyst deactivation and poisoning. So the operating cost should be high.
Is this rational to charge such a feed to the catalytic system directly or is it better to use the process to somehow get rid of metals at least? If we need to use the solvent deasphalting system at the upstream of two before-mentioned cases and draw off about 20% of feed as pitch, we will succeed to lower the operating cost and increase the reliability of catalytic system because of the elimination of the major part of the metals. But in the opposite side, we have missed 20% of primary feed as pitch that it is a low value product. So the profit margin of the residue upgrading cases will decrease. However, as a second question, can we miss 20% of feed charge at the expense of increment of catalyst life cycle?
26/06/2016 Q: How do we calculate the weight percent for NH4CL in the stream? (1)
02/06/2016 Q: Ours is Axens liacence CCR we are facing frequent problem of regen gas drier cyclomatic valve stucking and when valve opened some power type material observed must be desi ant. Does anyone suggest how to overcome this problem. Second issue is though regen drier working fine dew point varies from +5 to - 30 deg c. Is this because of activated alumina property loss or dis functioning. (2)
17/05/2016 Q: In our CCR we are facing with the feed side plugging of the Packinox CFE heat exchanger. Could you tell me the reasons of this phenomenon? Have anybody experienced the same? What about the solutions? Till when is it worth cleaning? (7)
13/05/2016 Q: How to limit the coke content on spent catalyst in a Platforming reactor? (8)
25/04/2016 Q: We are looking at alternative option(s) that could expedite the unloading of residue desulfurization unit catalyst (from reactors) other than typical vacuum-out/jack hammering approach.
We have heard about the CO2 explosive technique - and just wondering if anyone has any success stories with that?
Any other feasible approach to be explored?
09/04/2016 Q: I recently heard that passivation of Hydrotreating catalyst before shutdown will reduce the hazards by reducing the time of Inert entry in the reactors during shutdown. A process called Catnap passivation is used for the same. Can anyone brief whether it is useful and beneficial...? (1)
28/03/2016 Q: I am looking in the possibility of replacing the let down valves after the CHPS and HHPS bottoms (in Hydrocracker), with Power Recovery Turbines. Is there way to estimate how much power would I get?
And is it reliable to use the power recovery turbines instead of let down valves?
Other than that, are there other alternatives to recover that power?
22/02/2016 Q: Why is reactor feed introduced at top of the reactor? (2)
14/02/2016 Q: Is Tatoray reaction (Reactor in which C7+C9 gets converted to C6+C8 aromatics..licensed by UOP) prone to run away reaction? If it is then what are the possible scenarios when run away reaction may occur in Tatoray units? (1)
13/02/2016 Q: What are the conditions of auto-ignition of hydrogen rich gas, when leaking through any point of hydrogen recycle system to atmosphere? Is there any differences between hydrogen rich gas and hydrogen make-up gas regarding this issue? (1)
09/02/2016 Q: What will be the PONA content of vacuum gas oil which is used in hydrocracker? (1)
09/02/2016 Q: In hydrocracker unit what are the amount of heat of energy invloved for following reactions.
1) olefin saturaion
2) aromatic saturation
3) denitrification
4) desulfurisation
01/02/2016 Q: For how long we can run the hydrotreater in recycle mode , if less feed is available? (4)
20/01/2016 Q: How can I calculate heat of reactions in a hydrocracker unit? (1)
13/01/2016 Q: I m working in DHDS unit. where Reactor outlet material ( H2s, H2, Treated Diesel, N2 compounds etc) flows to series of Exchangers( 10 no.) tube side . Currently we found minor leak at drain line (CBD) of last exchanger outlet (press.44 kg Temp 125 °c) . Try to attend with online clamp with material filing( Benzola material used) but no effect. Any other material or method to fill up the leak? (2)
08/01/2016 Q: In our semi regenerative catalytic reforming unit we are facing problems of low recycle gas purity.70-72%. Octane has been in the range 92-93 as per design. Reformate yield is around 83% against design of 86.%.Our stabiliser off gas production has been incresed by 3-5 % and reflux control valve has been opened fully. Ratio of C1+C2/C3+C4 is higher 1.9. PCE dosing @1.5 ppm (8)
01/01/2016 Q: Our NHT feed is designed for C5-90 cut Straight naphtha with no olefins, We are facing problem of increase in reactor outlet temperature of approx. 5 degc(from 324 to 329 degC) for the last 2-3 days . our feed comes from storage tanks( N2 blanketing) which is directly coming from avu column
what could be the possible reasons for high temperature increment in NHT Reactor exotherm?
Please kindly suggest the causes
29/12/2015 Q: In our new octanizing unit (axens) we have gas dryer in catalyst regeneration section, but in some cases we need to bypass the dryer and we don't know can we do this without shutting down the regeneration section or not? (4)
25/12/2015 Q: We are getting very high iron and chloride content in our stripped water processed in sour water stripper unit which is used in our DHDS unit. What can be the possible reason and what we can do to minimize it? (5)
11/12/2015 Q: The bottom part of our atmospheric distillation tower operates at 350oC, and we are giving the superheated steam to the column at around 420-430oC. Let's say, we want to decrease the temperature of superheated steam to 400oC but still it is superheated steam of course. Might it have any effect on the distillation of the products or the quality of the products to decrease the temperature in this manner? (7)
04/12/2015 Q: We are operating HCGO/VGO as feed in Coker hydrotreater unit with system pressure of 95kscg. But actual feed quality is much better than design feed quality.In fact Hydrogen partial pressure is maintained more than design 91 instead of 86 with purity recycle gas as 95%.We thought of optimizing hydrogen consumption as well as power reduction by reducing system pressure . Can anybody tell ,will this attempt help us to bring power in MUG? otherwise how much can I go down further on system pressure ? And I wanted to calculate based on feed spec, how much gas oil ratio is required for my system? Can anyone share calculation basis to arrive gas oil ratio and system pressure to maintain? (3)
26/11/2015 Q: Recently, because of some difficulties, we have substituted demineralised water injection system with HP boiler feed water branched from HP BFW header. So, HP boiler feed water is being injected upstream of air cooler while it contains oxygen scavenger, amunium, and phosphate materials. In addition, the temperature of HP BFW is 80 degree centigrade higher than demineralised water's. By focusing on this, are there any consequences about this substitution for a long time of operation? (1)
19/10/2015 Q: We have Cyclemax continuous catalyst regeneration in CCR, but it has been observed that regeneration gas flow to Regenerator FI reading goes low and on low flow Hot shutdown occurs. Its LP and HP tappings have been cleaned , but still it shows low flow. Can you please suggest what else could be the reason and how to resolve such issue?? (3)
18/10/2015 Q: We are looking for a non hydrotreating based technology to decrease condensate sulfur content to lower than 200 ppm. There is a condensate stream in our refinery in which its sulfur content decreases from 3300 ppmw to 1000 ppmw by caustic wash and we need a further decrease of sulfur content to minus 200 ppm, but not with hydro desulfurization. Please advise. (5)
16/10/2015 Q: What is the max temperature up to which we can circulate recycle gas during start up without pre-wetting the catalyst in unionfining process to avoid sulfur leach out from catalyst? (Main catalyst KF 757 1.5E STARS) (2)
20/08/2015 Q: We have a two coke drum DCU. We are facing issue of high Sulphur content in our pet coke. The feed quality has been fairly consistent. Normally we have been operating with a cycle time of 18 hrs & COT of 501 Deg C since last two years. The issue of high sulphur has surfaced over the last two months only. Can anyone suggest the reason? (5)
19/08/2015 Q: We are operating our DCU with COT at 501 Deg C. The feed quality is also more or less constant. For the last month and a half months we have seen increased sulphur in our coke in excess of 9 wt %. Can anyone suggest the probable reasons?
10/08/2015 Q: In Hydrogen plant steam flow is fixed at 30-70 % plant load in Haldor Topsoe unit, while in Linde hydrogen plant steam flow is fixed in between 30-50%. What are the reasons for this difference? (1)
05/08/2015 Q: Could high moisture in recycle gas contribute to cracking in CCR reactor? (4)
03/08/2015 Q: Why is the Pitot tube placed at a 45 degree angle during installation in cooling water supply? What is the reason for placing at this angle?  
29/07/2015 Q: For VGO Hydrotreating units at the upstream of the FCC, we currently use COMO Catalyst. The unit is operated by maintaining constant Sulfur spec in the sweet VGO which goes to FCC. Typically VGO feed sulfur is ~22000wtPPM and Nitrogen is ~2200 wtPPM. After processing in hydrotreater, sweet VGO sulfur is ~1500wtPPM and Nitrogen is ~1000wtPPM. Now the question is that if I change the catalyst from COMO to NiMO catalyst and maintain remaining all operating parameters same, what would be the Nitrogen conversion if I operate the unit by maintaining Sulfur level 1500wtPPM which is same as earlier? Will the Nitrogen conversion increases because of NiMo catalyst or it remains same since we are constraining the unit severity by maintaining same sulfur level? (4)
13/07/2015 Q: What are the methods available to maintain the electrically traced liquid sulphur pipe (OSBL pipe) when the elbow leaks?  
10/07/2015 Q: We plan to purchase regenerated catalyst for our kero and LGO hydrotreater. We did regenerate our own catalyst, but never purchased one from an external company. What parameters should I pay attention to, what are the recommended limits for poisons and other parameters to guarantee a near-fresh activity and lifetime? (3)
29/06/2015 Q: Currently my plant is experiencing overhead vacuum fluctuation from 20 mmhg to 40 mmhg.
The design overhead vacuum is 20mmhg and maximum throughput is 20MB.
The ejector system consist of 3 stage ejector.
The first stage ejector consist of 2/3 ejector and 1/3 ejector load.
The second stage ejector consist of 3 ejector, and normally 2 out of 3 online.
The third stage ejector also consist of 3 ejector, and normally 2 out of 3 online.
We had perform field survey and found that the second stage ejector temperature is relatively low compared to the other ejector (26degC vs 70 degC)
Earlier, we suspect air ingress in to the ejector and we had perform online inspection. and indeed, we found 1 coin size leak at one of the first stage ejector and the leak had been repaired. however, the vacuum fluctuation is still there.
We had also verified all the other ejectors for leaks but unfortunately no leak was found.
We are also having issue with the ejector condenser. the third stage ejector outlet temp is relatively high compared to the other condenser (65degC vs 40 degC). This problem was there since a few years which had eliminate the condenser as the root cause of the fluctuation.
Currently we are trying to search of other weak point which can cause air ingress into the ejector/vacuum system.
Appreciate your feedback on the matter.

24/06/2015 Q: In cyclemax of of CCR unit, we had issues with fines collection for around 15 days. But thereafter when we performed the fines collection 350 kg of fines was removed. This surprisingly seems to be large amount. Is it due to new catalyst ? What should be the real amount of fines collection per day?
What could be the reasons for higher fines generation?
20/06/2015 Q: We are facing frequent issue of moisture carryover in the light naphtha stream to our hydrotreater. Does anybody have experience of putting a salt drier in the hydrotreater upstream to remove the moisture. What sort of damage it can cause to the catalyst? (5)
17/06/2015 Q: What are the causes of high pressure drop in Kerosene hydrotreater reactor (Kerosene hydrotreating Unit) and how can they be solved? (4)
16/06/2015 Q: In VGO HDS,fractionator Naphtha produced contains presence of H2S. So Naphtha is routing to slop. We have increased steam to stripper to reduce the H2S content in Naptha. But no effect obseverved.What can be done to avoid H2S presence of Naptha for routing to storage? (5)
16/06/2015 Q: We are observing high CS2 content in our straight run naphtha. This is not on regular basis but frequent and sometime it goes up to more than 20-25 ppm also. Please advise what can be source of such high CS2 content in naphtha intermittently.
The sources may be narrowed down to:
1. Presence of CS2 in Crude itself- Please suggest the probability of the same and if any known crude with high CS2?
2. Since CS2 formation requires very high temp, can it be formed in crude heaters? or any other process?
3. Though probability is less, can it come from recycle hydrotreated naphtha?
07/06/2015 Q: What is the difference between UOP R-234 and R-264 catalyst?
We have been provided later kind after first TA and replaced former.
07/06/2015 Q: In our CCR reactor, there is coking taking place due to which many a times L-valve assembly for catalyst regeneration line gets choked. As far I know Metal catalyzed coking is one of the reasons, so we maintained proper DMDS flow as per licensor. I would like to know what might be other reasons for coke lump formation in reactor. Coke lumps filling the scallops were found during TA opening.

Choking occurred in spent line and not regeneration. During TA coke lumps were removed from scallops by cutting and created window for coke removal and again welded, which unfortunately delayed TA.
03/06/2015 Q: Why is the pilot tube placed at a 45 degree angle during installation for provided reading in flowmeter of cooling water supply and return header? What is the reason for placing at this angle?  
30/05/2015 Q: We have hydrotherapy unit , consisting of cobalt molybedium (s-7 and s-120) reactor, reaction temperature 610 F system pressure 24 bar.We have a problem for two months that is the reflux drum of stripper got very low thickness observed, its boot water has PH 2.0----2.5, iron greater than 100 ppm while chloride was 1000 to 2000
condensate injection 8bbls/h from condenser inlet.
We have already done the cleaning of all heat exchanger , overhead condenser, overhead reflux drum.
Then start up of the unit was performed but condition remains the same.
Please share your opinion regarding this problem.
05/05/2015 Q: We are currently having problem in debutanizer of our naphtha hydrotreater due to ammonium chloride deposition. However, we do not have online water wash and we do not want to shutdown our unit. We are thinking of injecting steam (while the unit is commissioned). Is onstream injection of steam in the debutanizer to remove the ammonium chloride deposition applicable and effective in a debutanizer? If yes, what are the parameters we can check to safely conduct this activity? If no, are there any other way in order to remove the ammonium chloride deposition without shutting down the unit? (6)
25/04/2015 Q: Which is preferable for naphtha hydrotreater catalyst regeneration in situ or off situ , and if in situ what are requirements and steps for that process? (2)
23/04/2015 Q: I was asked for evaluation of doing the leak test and pressurization step in the Hydrocracking unit start up with Nitrogen instead using Hydrogen. We have 1 reciprocant compressor for make up and one centrifugal compressor for recycle gas, I would like to know what do I have to consider to make this evaluation, what I know by now is that my recycle gas molecular weight is 4 and N2 is 28, so my centrifugal compresor could not be able to increase the pressure more than 250 psi (aprox). what should I take in account?. Is there any gain doing this? (4)
21/04/2015 Q: I would like to ask about required H2/HC ratio and coker naphtha processing in a naphtha hydrotreater.
We have a unit processing a mix of straight run and coker light naptha. Unit consists of two reactors, one for diolefin saturation and one for HDS and olefin saturation, both use regenerated NiMo catalyst. Colleagues intend to raise coker naphtha ratio, which is currently maximized in 12%. I made some calculations which resulted, that coker naphtha has around 90 Nm3/m3 chemical H2 consumption, and the units H2/HC ratio is around 60-100 Nm3/m3 depending on throughput. 12% naphtha results in ~16 Nm3/m3 chemical H2 consumption. If I remember well, the H2/HC ratio should be at least 5 times the chemical consumption, in this case 5*16=80. Am I right, or can this value safely be reduced? Does anything else restrict the max ratio of coker naptha processing? Temperature raise is about 25-28 °C on HDS reactor with an inlet temperature of 290 °C.
13/04/2015 Q: I would like to know what is the difference between phenolic stripped wash water or non phenolics stripped wash water in the hydrotreater or hydrocraker process. (4)
25/03/2015 Q: I would like to know if when we design a transfer line of CDU or VDU heater then do we consider erosional velocity as a constraint? The mixed phase velocities in transfer line are frequently higher than calculated erosional velocity (from API-14E). (4)
23/02/2015 Q: We are trying to determine the appropriate lab test and normal analytical ranges in order to bring imported MVGO to our new hydrocracking unit. Licensor is concerned is about presence of Na and Cl but also other contaminants such as P. What are the normal ranges of Na, Cl and other metals to bring to the hydrocracking unit to avoid catalyst damage? (2)
21/01/2015 Q: What are criteria for providing the bypass line in the control valve? In NG fuel line from 32 ksc to 5 ksc pressure drop at 70 deg C, can we provide the control valve bypass? (2)
30/12/2014 Q: This question is related to kerosene merox unit. After processing kerosene in merox unit, what are the main reasons for poor saybolt color of kerosene product? If kerosene feed to the merox unit has saybolt color of +26, kerosene product from merox unit observes saybolt color of <16. Can someone explain the possible compounds which causes color problems to the kerosene product? If we go to Kerosene Hydrotreater, there will not be any issues of color problems and in fact it will be improved because of olefin and aromatic saturation. Please share any literature or chemistry related to the kerosene color problems in merox units. (3)
22/12/2014 Q: What happens to the catalyst if water goes to naphtha or diesel hydrotreater reactor along with feed which is having Nickel molybdenum catalyst. (2)
19/12/2014 Q: What are the major modifications are required for processing Light kero in DHDS unit in place of Diesel ? (4)
18/12/2014 Q: I'm doing a work about octane rating, but I haven't find the Research Octane Numbers of cis-1,2-dimethylcyclopentane and trans-1,2-dimethylcyclopentane. Who can tell me? (3)
19/11/2014 Q: What are the Pros & Cons in case of Hot start-up of Hydrogen generation unit? Why it is generally not preferred and also why is there no detailed procedure given in operating manual ? (1)
19/11/2014 Q: I am working in Hydrogen generation unit. Our naphtha vaporiser in HDS section got fouled frequently. What shall be the reason behind choking of naphtha vaporiser? (1)
14/11/2014 Q: My question is if we will do evacuation test for DHT reactors -20 PSIG How do I know how much offset of Positive?  
08/11/2014 Q: I am working in an HGU unit. I want to know if olefins or unsaturated compound increases, what will happen in prereformer catalyst.

01/11/2014 Q: in our naphtha hydrotreater with TK527(TOPSOE) CATALYST we have to treat a naphtha with 200 ppm oxygenate (MTBE), but we don't know is it harmful for catalyst or not??
operating condition:
feed=240 m3/hr
reactor inlet temp.=315 deg c
system press.=30 barg
(this unit is upstream of CCR UNIT)
03/10/2014 Q: In our DHDS unit we recently we noticed product colour changed after the reactors. After some days product colour reverted to normal. Can anybody please explain this colour changing mechanism? (5)
25/09/2014 Q: Can we process FCC's Clarified oil (CLO) or Decant oil as feed to Hydro cracker? My question is that Unconverted oil from Hydro cracker is usually good feed to FCC, So I would like to know if we process FCC CLO in hydrocracker then how much of it will it to convert to Unconverted oil in Hydrocracker? We will use filters to reduce catalyst content in CLO so that hydro treater won't get affected. (2)
05/09/2014 Q: In a hydrotreater I have normally seen shutdown valve provided on HC line exiting the HP separator. The shutdown valve will close upon sensing low level in the HP separator. Shutdown valve is provided between HP to LP interface. Level indication failure of HP separator will lead to gas break through from separator to stripper and stripper is not designed for such high pressure conditions.
But in a kerosene hydrotreater unit I have seen the shutdown valve tripping logic on high-high pressure in HC line exiting the HP separator instead of low level in HP separator ( this low level tripping is not present ). The tapping for pressure ( 2 out of 3 tripping logic ) is taken from downstream of angle valve (pressure reducing valve ). What can be the reason of changing the shutdown valve closing logic from low-low level in HP separator to high -high pressure in the HC line exiting the HP separator?
21/08/2014 Q: In a team discussion about the start up sequence of Naphtha Catalytic Reformer, everyone was wondering about the effect of prolonging the hot hydrogen circulation across the catalyst bed more than 12 hours.
The question rose from the fact that usually the Stabilizer (Debutanizer) Tower in the Reformer Unit is started parallel with reactor heating up, in such a way that when Reactor has reached the required temperature for charge in the liquid feed, the Stabilizer Tower has been ready to strip out the light ends. But it's not seldom that Stabilizer Tower is suffering from un-predicted prolonged problem --- such as very frequent bottom pumps' strainer blockage-- while reactor inlet temperature has reached the feed cut in temperature.
Under such situation, the start up team was in the pro-cons whether to keep the hot hydrogen circulating across the reactor for few more hours till the readiness of Stabilizer Tower, or to immediately cool down the reactor loop. The first opinion merely consider about the time efficiency, while the second group worried that alumina support of the catalyst will undergo a phase change due to hydrogen embrittlement on this alumina.
Has anyone here had the similar experience, and can give us more enlightenment on this matter?
19/08/2014 Q: I'm currently working on a VGO hydrocracker simulation. I want to know some common problems in normal industrial operation in this kind of process. (1)
19/08/2014 Q: Retained sample of our Naphtha product found to be decoloured after a couple of months. Naphtha product was w a mix of Straight run and cracked and hydrotreated naphtha. Is there any particular reason for colour degradation of Naphtha on storing? (1)
31/07/2014 Q: What is the limit of Oxygen content in Naphtha feed for Hydrotreater to avoid gum formation?
What is the ASTM method to test Oxygen in Lab?
31/07/2014 Q: Isomerization plant Molex unit has been commissioned.How to asses the Adsorption capability of Adsorbent.
Similarly how do we know that Rotary valve parameters are fine tuned ? Separation efficiency would be affected if any of the two above mentioned factors are not optimized.
24/07/2014 Q: I am working in Hydrogen generation unit. I want to know whether if naphtha preheater tubes got a leak and super heated HP steam went to naphtha side then would superheated HP steam go to hydrogenerator (Co-Mo catalyst)? What is the effect of steam on Co-Mo catalyst life? (4)
17/07/2014 Q: Is there a way we can assess VGO Hydrotreater is running at Aromatic saturation equilibrium or still scope to increase Aromatic saturation by increasing catalyst volume? We are looking for revamp options of existing unit. Since operating pressure cant be changed to increase aromatic saturation, the only option is to add catalyst by introducing new reactor bed. The question and what I need to prove is that increasing catalyst volume improves the Aromatic saturation. This can be done by proving the current unit is not running at Equilibrium controlled all the times. During EOR conditions, it may reach equilibrium controlled because of high WABT. I need some literature on this. Can anyone share your thoughts? (5)
11/07/2014 Q: In our Hydrogen Generation Unit HP steam silica level is running high at about 0.1 PPM against design value of 0.045 PPM. We maintain BFW Ph-9.5, excess Phosphate - 3 PPM, Hydrazine excess 0.1 PPM and continuous blow down Gestra valve is 100% open. Conductivity and TDS is normal. How can we reduce silica? There are no BFW exchanger leakages
24/06/2014 Q: How to calculate amine requirement to absorb H2S in hydro treatment unit by just knowing the feed sulfur content?  
24/06/2014 Q: I am working in Heavy cooker gas hydro treating unit.In plant,RGC primary seal drain is connected to CBD.There are chance that by mistake or passing of isolation valves, CBD gets pressurized with 100 kscg gas.Kindly tell me the reason why high pressure drain is connected to CBD?Is not it possible to use pressure reducing valve as safety measure in that line.  
26/04/2014 Q: In the Design of a DHDT unit, what are the criterion for selection of a HHPS (Hot High Press. Separator), along with a CHPS (Cold High Press. Separator) or with CHPS only unit?
Also, in some configurations of DHDT, we can see a HHPS + CHPS + CLPS (Cold Low Press. Separator / i.e. Flash Drum)
What are the criterion for such selection ? Also what are the advantages and disadvantages of the 3 options?
26/04/2014 Q: In mild hydro cracker, we have amine absorber column to reduce/remove H2S from Recycle gas. Due to amine foaming, amine is carry over to recycle gas compressor.
Any one can share their experience to reduce foaming of amine or how to do Oil skimming in amine columns and what are the parameters to be monitored while doing this activity.
How can we know Oil skimming is completed and there is no foaming in the column?
21/04/2014 Q: Mild hydro cracker unit fractionator furnace pass flow Flow transmitter impulse tubes are cracked due to Header vibrations. We faced this kind of problem in case of unit is upset due to recycle gas compressor failure. what are the possible chances to get multiple phases in this header.
Back ground: Product stripper outlet ....>G-0006...> Exchanger E0004 (Shell side - Pump out let liquid, tube side Reactor effluent) ......> Exchanger E0025 (Shell side- E0004 O/L, Tube side fractionator O/L)
This incident was happened after start up of pump G-0006.
Any one experienced this kind of problem?

19/04/2014 Q: We have an Axens CCR unit with 4 reactors and I would like to know whether we can maintain different temperature in each reactor? If the answer YES, what are the effect on RON and catalyst? Please share your experience. (2)
13/04/2014 Q: We have a Steam methane reformer having side fired self respiratory burners. To attain the correct O2 in flue gas of primary reformer, burner dampers are being adjusted. What is the correct sequence for throttling the burners? Should the bottom most burners should be throttled more than the top ones or vice versa?  
11/04/2014 Q: In case of side fired self respiratory burners in reformers what is the correct sequence of adjusting the air?
From bottom row burners to top row burners in increasing trends:
in 1st row 40%, 2nd row 40%, 3rd row 40%, 4th row 30%, 5th row 30% & 6th row 30%
in 1st row 30%, 2nd row 30%, 3rd row 30%, 4th row 40%, 5th row 40% & 6th row 40%.
This flue gas is going to convection section for heat recovery.
29/03/2014 Q: We are having Feed Surge Drum in Diesel Hydrotreating Unit, for maintaining pressure of FSD we provided Blanketing Hydrogen and relief to LP Flare. Fail safe positions for the Control Valves in Hydrogen is Fail Open, LP Flare is Fail Close (Where as it was reverse in previous company where I worked last). If in case of Air failure Hydrogen to FSD CV gets open and may get pressurise as there will be no any relief
What may be the basis of selecting the fail safe position of both CVs?
21/02/2014 Q: In the crude distillation unit, we face problem with Gas Oil colour. Any one have any idea to solve this problem or any one have seen like this in any refinery?! (8)
05/02/2014 Q: Is there any SO2 production due to decomposition of Sulfolane use as a solvent for aromatics extraction? (2)
04/02/2014 Q: We are trying to figure out how to improve the feed control to our new Hydrocracking and Hydrotreater Units, since one of the feeds comes from the Coker Unit, we want to know how variable are the quality and flow of the HCGO, Naphtha and LCGO, because we are aware it would be changing while coker cycles are taking place. We don't have tanks to store LCGO and Naphtha as feed to the units, so these streams go to the hydrocracker and hydrotreater directly from the coker stripper, and if there is a sudden change in composition or flow, it could lead on a runaway. (3)
20/01/2014 Q: I am currently working in diesel hydrotreater plant. The end products are naptha, kerosene and diesel.
According to the lab reports the sulphur content in naptha is 2.7ppm and that in kerosene is 0.5ppm. Kerosene being a higher molecular weight fraction should have higher sulphur content. What is the correct explanation for this?
13/01/2014 Q: We have steam methane reformer. The outer surface of the tube is having deposits and leading to high fuel consumption as well as high temperature in flue gas side in waste heat section. During turn around we want to clean the outer surface of reformer catalyst tube so that we can reduce the fuel consumption and reduction in waste heat section temperatures. Is there any standard method available to clean the outer surface of the tubes? (5)
08/01/2014 Q: Lately, have been experienced tube leak in DHDS stripper feed-effluent exchanger, Tubes were plugged and hydro-tested.
Four months later, again leak developed and found tubes in bad condition, and was recommended for full bundle re-tubing.
I would like to know what could be root cause for this tube failure in short time? Any specific improvement need to be done on internals of exchanger?
12/11/2013 Q: In DHDT unit, RGC anti-surge control valve is opened at around 13-15 %, at this opening deviation from surge line is 0.14-0.20. The design molecular weight of Recycle gas considered is 2.94 whereas actual Recycle gas molecular weight is in the range of 2.25-2.4. Suction temp/ Pres:63 deg C /110 kg/cm2;Discharge temp/ Pres:90 deg C /131 kg/cm2
1.0 Can we fully close the anti-surge valve in order to increase energy efficiency of RGC ?
2.0 what other actions can be taken to minimize RGC anti-surge opening ?
3.0 By Incorporating the actual recycle gas molecular weight in anti surge controller block and compression suction flow transmitter, will there be any improvement in deviation from surge line ?
09/11/2013 Q: How H2/HC ratio is calculated? Our design value(in VGO Hydrocracker) is 843 at hydrotreater inlet .Is it just a ratio of H2 gas flow to VGO flow at hydrotreater inlet mixing point.
Shouldn't there be any factor of purity of recycle gas to be incorporated?
Is H2 gas flow & VGO flow should be temp. corrected values or it is just a ratio of the tags showing in DCS without temp. correction.
02/10/2013 Q: Does anyone know where I can get access to a bauxite percolation pilot plant or contract manufacturer to decolorize wax?  
23/09/2013 Q: Can we change the trash baskets in heavy naphtha hydrotreater reactor by ceramic balls?
(in all turnarounds trash baskets are clean without any scales but assembly of 85 baskets is very hard!!!)
catalyst volume is 30 cubic meter
20/09/2013 Q: What is the maximum cracked feed (LCO+LCGO) percentage that can be processed in the DHU unit? (5)
18/09/2013 Q: How can I calculate the optimal velocity in furnace tubing? At our gasoil/kero hydrotreater we operate usually at low throughput, but we keep the recycle gas at a higher value than needed for the reaction, to prevent the coking of furnace tubes. I guess that the optimal recycle gas amount could be calculated, but I don't know how to do it.

Some additional info: It's the unit manager's explanation that he doesn't want to decrease recycle gas to prevent heater coking. We are usually running on low throughput with 4-500 Nm3/m3 H2/CH ratio. In the last cycle we had pressure drop problems on our reactor, we found solid deposit on top of the bed. We performed a furnace coke burning process during the last turnaround, and found that there was some significant coking in the furnace. Our licensors suggestion is, that H2/CH ratio should be approx. 5 times the H2 consumption. Based on this, 100-200 Nm3/m3 would be enough, but we are running often at 400-500 ratio, which is way higher than suggested.
09/09/2013 Q: I am process eng. at diesel hydrotreater unit. Recently we decided to treat blending naphtha in diesel unit
spec of blending naphtha: IBP=145 dry point=195, sulphur in feed= 2000 ppm, sp.gr of feed is= .7640
flash point=35 deg c. We are planning to produce solvent aw402 from blending naphtha but we have two questions:
1) catalyst volume is 75 m3 and we are worried about LHSV (design capacity for diesel is 18000 bbl/day and minimum throughput of unit is 60%in design case).
Is minimum 200 m3/hr for new feed OK or we can process lower feed?
2) What's the minimum inlet reactor temperature to hold minimum cracking because we have problem to set flash point in stripping section. Can we reduce temp. below 290 deg c ?
21/08/2013 Q: What can be the cause of coloration (yellowish green) in VGO Raffinate hydrotreater effluent?
13/07/2013 Q: If Light Refomate is the only feed of an Isomerisation unit, is there any chance of reflux drum of stabilizer not getting liquid at normal working pressure of 13.5 Kg/cm2g? (Liquid level building at lowering the pressure to 11.0 Kg/cm2g) Is it because of the C4 content is more in reformate? (1)
18/06/2013 Q: 1. VGOHT is processing Hot Feed from DCU & LVGO/HVGO/HGO from CDU-VDU and combined Crack + Straight Run components from Feed Stroge Tank. I would like to know about quick inference / thumb rule to access %Crack Components in Feed from Sorage Tank by monitoring Reactor Profile and Hydrogen Demand. (2)
14/06/2013 Q: We have in our plant diesel hydrotreater unit Packinox Exchangers high dP increment. The dP of Packinox feed gradually increased from 10~12 psi and reached the low limit alarms which is 22 psi in last one month duration. Preliminary observation is showing that the raise in pressure occurred after the increasing of SC#6 from 11 MBD to 15 MBD
What is causing this problem, and what is the solution?
10/06/2013 Q: We planned to carry out top layer catalyst skimming in our Naphtha Hydro-Desulfurization Treatment (NHT) reactor. The skimming amount will be about 25% of total reactor volume. For this purpose, we have purchased fresh UN-SULFIDED Co-Mo catalyst, that will be loaded on the top of old sulfided catalyst during the near future skimming activities.
In order to anticipate the future unit re-start up with PARTIALLY UN-SULFIDED catalyst, we consulted the catalyst manufacturer how to carry out the IN-SITU PRESULFIDING for this new unsulfided catalyst with the presence of old sulfided catalyst underneath . But the recommendations were not convincing.
** The presulfiding will be "liquid phase presulfiding" where the Sweet ( Treated ) Naphtha is circulated through the NHT reactor under H2 environment, before DMDS injection commences at reactor temperature of 180 degC. The required amount of DMDS for new unsulfided catalyst will be injected at the rate of 0.2%-wt-S of the circulating Naphtha. During this DMDS injection, reactor inlet temperature will be raised up gradually to 270 degC where the H2S breakthrough will happen, and will be on hold at this temperature till all required DMDS is injected. The Unit will then be adjusted to get on specification Stabilizer bottom product before being put on once through operation**
Please advise regarding to below questions :
1. Will the old catalyst --which had been sulfided in the past -- get the adverse effect during this future presulfiding, such as washed out sulfur from old catalyst surface ?
2. With the presence of old sulfided catalyst under the new un-sulfided one, we are not sure whether the measured exotherm across the reactor, and also the H2S breakthrough ,will be representing the actual presulfiding progress. Because old catalyst will enhance the exotherm and will advance the DMDS breakdown into H2S.
In order to protect the new catalyst from thermal damage, what can we do to minimize the exotherm effect from the old catalyst to the new one ?
In order to determine the end of presulfiding, can we rely on the total amount of required DMDS that has been injected during the presulfiding ?
04/06/2013 Q: In Naphtha cracker the molecular sieve dryers are regenerated with Methane. But in Isomerization plant(MSQU) the molecular sieve dryers are regenerated with the valuable product isomearte. Is it possible to regenerate with off gas or fuel? is any limitation as these streams contains sulphur...? (1)
01/06/2013 Q: What is the purpose of MFA additive in MS blend? What is chemical composition of MFA? (1)
31/05/2013 Q: On line Hygrometers are used to measure moisture content in CRU recycle gas. On many occasions the measured values are suspected to be incorrect by plant people. What type of on line hygrometers are highly reliable and what are the maintenance practices for these analyzers? (2)
22/05/2013 Q: What is the purpose of adding Ammonia during the regeneration of HDS reactor for Gasoline? (3)
18/05/2013 Q: Is it possible to Hydrotreat Merox treated Naphtha in NHT? (4)
12/05/2013 Q: My question is on Acetylene Selective Hydrogenation Catalyst (Palladium –Pd based with promoters):
Ethane gas gets cracked in the Cracking Furnaces and the effluent goes through series of processes that includes quenching, heavy contaminants / heavy hydrocarbons removals, Multi-stage Compression, Caustic Scrubbing with Drying leading to De-Ethaniser (DeC2), and DeC2 Column Overhead vapour to the Two-stage Acetylene Hydrogenation Reactors. Main feed Ethane gas has a spec. of CO2: 200 to 1000 ppm; Total Sulfur: 500 ppm; Moisture content: 100ppm and it is directly cracked in the Furnaces. There are other feed streams having Sulfur ppm in the range upto 50 or so, with metal traces at lower ppb levels. The Reactors are operated with Carbon Monoxide level of 1000 ppm to 3000 ppm Max or so, at the upset conditions. Outlet Acetylene ppm levels are stringent in the range of 0.2 to 0.3 to produce Ethylene with 1 ppm Max Acetylene impurity.
a) Pl. let me know what all process parameters have direct impact on Catalyst deactivation and thereby short run-time requiring ex-situ Regeneration.
b) How will you control the parameters effectively to have much longer Catalyst run-time?
c) What is normal catalyst run-time for such Catalysts irrespective of any Catalyst vendors?
d) Whether going for Regeneration, would it be recommended to revive activity and selectivity to that of fresh material? Any risk involved in taking decision in favour of Regeneration?
e) Vendors confuse often with jargons, Reactivation and Regeneration. Are they one and the same or the process of reviving the spent material to the active phase to prolong the operation with recycle not only due to downtime of plant but also, expensive nature of catalyst with precious metals?
f) Pl. suggest suitable catalyst vendors with whom development activity can be collaborated with the company’s R&D Centre.
g) Any other important points in relation to specific Catalyst poisons, improving run-time atleast upto 4-5 years if not 10 years+

Your thoughts on this, in whole or part, greatly appreciated.
19/04/2013 Q: In the case of multistage compressor, why is it that the compression ratio are not set equally?  
14/04/2013 Q: We have proposed to install a "hot separator" in the recycle gas loop of a heavy naphtha hydrotreater unit operating at a pressure of 65 barg. The configuration which we are contemplating will be similar to kerosene hydrotreater units which usually have both high temperature and low temperature high pressure separators to enable liquid condensed at high temperature to be directly sent to the stripper without cooling. Concerns have been raised by the process licensor that chlorides may be condensed with liquid in hot separator and reach the stripper with feed thereby causing choking/corrosion in stripper. Does anyone have any experience of operating a hot separator in a heavy naphtha hydrotreater? Simulations predict significant heat recovery potential from this project in line with savings achieved in several similar projects in Diesel Hydrotreater units. Can anyone share experience of discovering precipitated chloride salts in the high temperature combined feed effluent exchangers (operating over 200 deg C) during cleaning operation?  
31/03/2013 Q: What is the typical level of diolefins that corresponds to Existent Gum and Potential Gum Specificaions for Light naphtha and Heavy naphtha? (1)
29/03/2013 Q: Please give the possible causes of increased pressure drop in middle and lower catalyst beds in VGO Hydrotreater main reactor. What solutions could be implemented to prevent pressure drop events?
08/03/2013 Q: In VGO hydrocracker and Hydrotreater, where would you put the purges for the PSVs on CHPS, HHPS etc to prevent it from plugging/fouling due to salt deposition i.e. on the common take off from the vessel or separate for each PSV, right below the seat of the PSV? (1)
20/02/2013 Q: We have NHDT for Isomerization unit. In NHDT wash water used to dissolve the salt, thereafter sour water separated from naphtha in separator vessel. For last 2 months we are unable to separate sour water in separator or even in stripper column. What could be reason and solution? (2)
29/01/2013 Q: is it possible to Presulphide NHT Catalyst with Straight Run Diesel ? Straight run from CDU can also contain moisture. Can this moisture result in lump formation or reduce the strength of the catalyst?  
28/01/2013 Q: 1. How we predict/calculate theoretical life of sulfur guard bed if we know the weight and bulk density of catalyst? Purpose is to absorb sulfur (Organic/inorganic H2S) from light naptha before goind to isom section?
2. Can we use DIH (De-ioshexaniser column) bottom isomerated in Recycle feed in Isomerisation section for increasing the conversion of n-parafins into iso-parafiins? what will be the impact on isom catalyst life?

Additional info:
Here light naptha means C5-90 CUT coming from Naptha splitter (Ex-cdu unit)
26/01/2013 Q: Planning a New Hydrotreater (HDS) Unit to remove Sulfur from Heavy Cracked Naphtha stream ex.FCC Unit. To achieve Ultra-low Sulfur Gasoline specs. (<10ppm), and also to meet Gasoline production, other two Naphtha streams are planned to be treated in the above new Hydrotreater. Naphtha ex. Merox Unit having about 375ppm Sulfur is one of them. My query to subject matter experts are as under:
a) Is it recommended to treat Merox treated Naphtha in the above new Hydrotreater? OR
b) Bypass the Merox Unit; Only treat the Naphtha stream with aq. caustic (Pre-wash only for H2S removel); No conversion to Disulfides by 2nd stage Oxidation process in the existing Merox Unit. Feed the untreated Naphtha stream directly to the New Hydrotreater.
Would appreciate pros and cons of treating the straight run-Naphtha, by-passing Merox Unit and treating it in the New HDS Unit. Many thanks for any useful advise from SME.
24/01/2013 Q: Recycle gas from high pressure separator is being treated in amine absorber. Treated gas then routed to recycle gas compressor (2 stage). We are observing lot of liquid accumulation in the inter stage KOD (after cooler). Liquid sample is analyzed and found that 99.7 vol% of sample contains water and remaining as amine. We are adding water to the amine system to compensate the loss through inter stage KOD.
Lean amine to absorber temperature is being varied in the range of 55 to 60 °C to maintain design DT between amine & gas. We have observed that the absorber top temperature is high when compared to bottom temperature.
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 0.3 to 0.5 wt% only.
What is the ideal delta temperature to be maintained between gas and amine?

18/01/2013 Q: Is it possible to Presulphide the NHT catalyst with Fuel gas containing H2S? (3)
10/01/2013 Q: Im working in diesel hydro desulphurization unit. We have one stripper-stabilizer section with common O/H system. Direct stripping steam is used here. We used to get product diesel Cu strip result as 1b. Recently we are not able to meet product copper strip of 1 (we are getting it as >3).
The following are attempted to normalize the section:
a. Checked the steam line for condensation. Line temperature is at 300 °C and there is no water
b. Upstream & downstream exchangers were checked for leaks. No leaks
c. Inlet temperature increased gradually from 235 to 254 °C
d. Stripping steam increased gradually from 3 MT to 4.5 MT
e. Withdrawn more distillate from reflux drum.
After doing the above, we could reduce the result from > 3 to 2.
Simulated the column with earlier conditions & with present conditions. 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?
2. How to find out whether there is flooding in the column. DP across the column is at 0.25 Kg/cm2. Will it vary severely if column is flooding?
06/12/2012 Q: Is jet flooding and or foaming a composition only issue for a hydrotreater stripper column? (3)
18/10/2012 Q: We are operating a hydrotreating reactor loaded with topsoe catalyst TK-562 Brim.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 1.5 bar. 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. We are operating our unit at 125 KBPSD. What is causing this? (4)
11/10/2012 Q: Recently we faced an emergency in our one through hydro cracking unit. 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. The logic behind introducing the emergency coil steam with pressure of 15 kg/cm 2 was if air will ingress through ruptured tube and make its way to vacuum column, it will lead to explosion of column. 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?
01/10/2012 Q: We are using light isomerate (c5/c6) for regeneration of Feed Dryers in Msqu Unit
As moisture is poison for isomeraisation catalyst so why is steam used for heating isomerate initially in vaporiser (shell side-naptha, tube side --steam) from 50 dgec to 147 degc (heating to the boiling point of isomerate)?
Then superheater is used for heating isomerate outlet temp. from 147degc to 320 degc. @ 60 degc /hr
1. If vaporiser tube punctured/ leaks then there may be chances of going steam into isomerate. Is process disturbed and will it poison the catalyst? (steam pressure~ 8.5 ksc while isomerate press.~9.0)
2. Why can't we use superheater initially from 50 degc to 320 degc? Is there any problem of heating liquid isomerate or it can lead to coke?
3. Is superheater used for only vapor phase not liquid phase?

30/09/2012 Q: I want to boost the pressure of MP separator off gas containing 90% H2 from 24 barg to 70 barg. 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? (1)
27/09/2012 Q: We have a problem in the hydrotreated filters when feeding HCGO. In these filters, we usually feed VGO and we haven´t any problems, but when we try feed HCGO from coker unit, the filters are plugged at the few minutes.
The ratio between HCGO and VGO is 30/70% aprox. and the temperature of these filters is 170ºC.
The filter element are wedge wire with 75 microns.
When the filter is plugged, although the filters are backwashed, the AP don´t go down, and It´s necessary to shut down the unit to clean up the filters mechanically.
We are doing some studies to identify the origin of the problem.
- Filtration studies to cuantify the solids of both of feeds: HCGO has 150-500 ppm of solids, which are mayoritary coke. VGO has 300 ppm of solids, which are mayoritary inorganics particles.
- Asphaltene determination (IFP method): HCGO has 200-500 ppm and VGO has 100-300 ppm.
- Compatibility studies: We have done a compatibility study in laboratory, which consists of adding gradually HCGO to VGO, then the mix is viewed in the optical microscope to identify the asphaltene precipitation. In this study we have seen that the feeds are unstable above 15-30% in function of temperature. The higher temperature the higher unstable is the mix, and the asphaltene precipitate at lower HCGO percentage.
Therefore, we think the plugging problems are due to the precipitation of asphaltene forming an impermeable layer on the filter, which doesn´t disappear even when the filters are backwashed.
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.
10/09/2012 Q: How does Vapor/ Liquid ratio at the bottom tray and Reflux to Feed ratio affect stripping quality? What happens when number of trays is increased? (1)
10/09/2012 Q: Why do Olefins have a higher Cetane Number than I-Paraffins? (2)
02/08/2012 Q: Is there any possibility of runaway reaction in Naphtha hydrotreater? the feed is contains olefins 4 wt% and sulfur around max 2000 ppm. (2)
17/07/2012 Q: What is the main difference between Accumulation and Over pressure for a relief valve? could anyone explain their importance while sizing a relief valve. (1)
17/07/2012 Q: What is the significance of PSV's discharge coefficient? how it will impact on relief valve sizing?  
13/07/2012 Q: Currently our client makes all jet A-1 fuel from a conventional Merox treating process. 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.
Quality assurance/control team has indicated that the jet product needs to be within 0.5 API of each other in the top/middle and bottom sections of the tank. To achieve this tank mixers or a re-circulation system is being considered. Based on API 2003, section 4.5.5 indicates that conventional low-speed propeller mixing has been in use for many years without evidence of problems from static generation.
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?
2) Do they use a nitrogen blanket system as a safeguard to protect against static buildup in the tank when mixers are applied?
3) Is it true that as long as the jet liquid level in the tank does not drop below the mixer elevation in the tank that there should be no concerns with static buildup in the tank? (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.)
4) Is the only concern on start-up or shutdown when a tank is being de-inventoried for maintenance or initially filled?
12/07/2012 Q: In our refinery in India after revamping of Hydrocraker unit for increase of Hydrocraker unit load to by 30% than earlier, Fractionator column perfomance seems not stabilize at all. There was minor modification done in fractionator column for increase of 30% load by replacing some earlier tray with high capacity valve trays & providing some packing ring inside column. But we are now facing problem of high column pressure of around 1.45 Kg/cm2 against design of 1.1 Kg/cm2. 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. To compensate for flash point & removal of naphtha from product we always kept column top temeprature at higher side of 99 0c than design 93 0c. 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 @360 0C than design of 374 0c. Presently we bring down the column pressure to nearly design pressure of 1.1 kg/cm2 by running two offgas compressor. 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 374 0c. (3)
29/06/2012 Q: 1. On what design basis would a vendor recommend pilot operated safety valves in a refinery? In our MSQU unit, they are used in RGC and Make-up gas compressor whereas in HDT unit they are used in H.P. separator and Stripper.
2. How are they different from normal spring type PSV?
27/06/2012 Q: I'm a process engineer at a CCR unit. Recently we have huge amount af gas formation because of extreme cracking, and RONC is 97 with total delta t equal to 218 deg c,WAIT=509, EDC IS 0.9 LIT/HR without water injection and our water content in platrecycle gas is out of service! I guess there is a chloride/water imbalance, but I want to reduce cracking reactions to increase RONC so can I have your recomendations please? (2)
17/06/2012 Q: Excuse my ignorance; I am neither a chemist, nor an engineer. My question revolves around RVP and RON properties of aklylate and refomrate.
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.
Questions that are driving me nuts:
1) How can a naphtha feed (with more saturated content) going into a catalytic reformation can have a higher RVP, than the resulting reformate which has more aromatics (unsaturated, cyclical = more volitile) content?
2) Why does alkylate which has more saturated molecules than reformate have lower RON and higher RVP?
3) Which is the most important contributor to RVP of a gasoline blendstock - the length of the hydrocarbon chain or it being un/saturation with hydrogen atoms? In other words, which has higher RVP - an unsaturated aromatic benzene molecule or a saturated paraffin pentane?
17/06/2012 Q: Excuse my ignorance, I am neither a chemist nor an engineer. I am actually in finance trying to figure out product properties.
I am having hard time understanding how is it possible to have a >10psi RVP naphtha feed go into a catalytic reformer, and have a <4psi RVP reformate come out the other end?
My logic dictates the following:
1) The lighter the molecule (shorted hydrocarbon chain) the higher the RVP and lower the Octane (and vice versa)
2) Unsaturated molecules are more reactive and therefore will have higher RVP compared to saturated chains of equal carbon atoms.
The questions driving me nuts:
1) Why does alkylate which has more saturated content have higher RVP than reformate?
2) Is the reason for akly having lower RON than reformate the fact that its molecules are lighter than in reformate?
19/04/2012 Q: MEG regeneration system. In our plant we have 2 rich MEG tanks that receive MEG/Condensate/Water solution from condensate flash vessel. Last time during pigging activities we receive many sludge from offshore, and now all this sludge is settled down inside Rich MEG tanks. MEG Regeneration package performance rapidly reduced, pumps could not deliver Rich MEG to regen, strainers getting clogged very fast, HC compartment of MEG flash vessel in MEG Regen package filling rapidly. 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.  
16/03/2012 Q: if we want to reduce the Hydrogen purity in DIESEL HYDROTREATER the current H2 purity is 99.99 fro PSA Unit now we want to take from other plant (Rheniformer units),During hydrogen plant turnaround, PSA’s are not in operation and only the off gas from Rheniformer units, low purity hydrogen, is available. 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.
the hydrogen from PSA :
HYDROGEN ------> 99.99
C1------> 0.1
CO + CO2 -----> 20 MAX
H2S ----> ZERO
HCL < 1
***** new hydrogen make-up with a reduced purity, coming from Rheniformer units
H2 87.5
C1 5.8
C2 3.6
C3 1.6
C4 0.4
C5+ 1.1
- what is the side effect of low purity for all the plants, recycle compressor, make up compressor and the load in addition the make up it is suitable for for such low purity?
What is the impact on
-Reaction section
-Product quality
-Recycle compressor?

05/03/2012 Q: I am currently exploring the possibility of selling Slurry as Carbon Black Feedstock. Although most of the expected qualities of slurry are able to meet the specs required of the Carbon Black Feedstock, the slurry is still high in CCR (~20 wt% in Max LPG mode) versus the required spec of < 10 wt%. For an RFCC, is there any operational adjustment that can be done to meet the CCR specs? (4)
20/01/2012 Q: I am currently managing a high pressure water injection triplex pump in a hydro cracking unit. 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. This particular job was given to my company on short notice and the only information i have received is that this was critical that the unit still perform at at least 50% production and in the event of a failure of the pump I'm operating the best thing I can do is run. 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. (3)
02/01/2012 Q: Naptha generated from Hydrocracker is having less then 5 ppm sulfur mainly of marcaptan type and the CRU feed demands for a sulfur level of less then 0.5 ppm. Existing NHDTs are having T'put limitations. Please recommend some other alternative for the removal of the marcaptan sulfurs so that it can be directly routed to the CRU feed pool. (3)
22/12/2011 Q: I am operating Pre-topping column at 2.4 kg/cm2 and 118 C . M.P. steam at 12 kg/cm2 at 550 kg/hr is fed as stripping steam. Gasoline reflux at 28 m3/hr and 13m3/hr of gasoline is taken out of overhead condenser. How do I calculate the dew point of overhead vapors? (3)
21/12/2011 Q: My question is related with high sulphur content in LPG from crude distillation.
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. The high S content is mainly due to high dimethylsulphide (DMS) and dimethyldisulphide (DMDS). We measure high DMS and DMDS both at the entry and outlet of the amine absorber and LPG Merox. 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 (??)
My questions are:
- What could be the origin of DMS and DMDS (synthetic / heavy crudes, slops processing...)?.
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?
- If DMDS is in the crude, according to its boiling point it should end in the heavy light / heavy naphtha. Has anyone observe high DMDS in LPG in his refinery?
- Could DMS and DMDS increase or decrease in the amine plant or Merox? (I do not think so). Are these compunds partially soluble in NaOH or could be removed in the sand filter?
- DMDS could also come from the circulating NaOH in Merox plant, if quality is not good (high concentratrion of disulphide in NaOH)? What is the normal or recomended concentration of disulphides in regenerated NaOH?
- What could be the alternatives to remove these compounds? I expect that nothing can be done in amine / Merox, these compounds are not reactive.
13/12/2011 Q: My question is related with a problem of copper corrosion strip failure (ASTM-D130) in gasoline. We have two tanks of off-spec gasoline:
- Copper strip corrosion 3B; SH2=0ppm, mercaptans = 9ppm. Does not improve copper strip corrosion test adding corrosion inhibitor
- Copper strip corrosion 2C; SH2=0ppm, mercaptans = 5ppm. 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?
- Could a NaOH carryover from the Merox unit cause a failure in copper corrosion test?
- Any other sulfur compound, besides SH2 and mercaptans, could cause this copper strip corrosion test failure?
- Does anyone know any commercial additive for mercaptan removal that could be useful for this problem?
10/12/2011 Q: Our Vacuum gas oil hydrotreater unit is operating at 125 KBPSD. We are processing feed, HAGO, HVGO. LVGO from crude unit and HCGO from Coker unit.We are having 24 filter for removing contaminants including one backwash filter.Our filter dp during steady state operation is 0.7 to 1.0 bar. We often face problem of high pressure drop of 3.5 to 4 barduring crude unit HVGO pump changeover or during taking of HVGO exchanger in line.In crude unit there are 2 HVGO pump one running and other standby with total flow rated capacity of 1770 M3/Hr.Running capacity 1650 including pump around,IR and product HVGO to our unit.Normally we consume approx 440 to 470 M3/Hr HVGO during steady state. I am not able to find out the root cause for high dp across feed filter during such activity in crude. This results in throughput reduction in ZVGOHT unit. Please suggest the possible cause for increase in filter dp (3)
10/12/2011 Q: I am currently working in Naphtha hydro treater, In our naphtha hydro treater stripper and Splitter both column is there, stripper is for removal of h2s and splitter is for removal light naphtha. 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.
07/12/2011 Q: We are operating vacuum gas oil hydrotreater unit after shutdown at 125 KBPSD. At 125 KBPSD we checked the vibration of Combined feed heater connected lines and tubes. 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 121 KBPSd. But still vibration persists. Can anybody guide me about sudden increase in vibration? (1)
06/12/2011 Q: Could you please guide me on ther Naphtha-expansion calculation i.e. the calculation procedure for naphtha, liquid to vapour at various temperatures  
06/12/2011 Q: In my refinery we have Diesel Hydrotreater unit which has undergone an emergency shutdown, following the tripping of recycle gas compressor. 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 1.5 vs normal 1.0 MMSCFH and lower system pressure of 620 vs normal of 670 PSIG was observed.
My question is where is go the higher hydrogen consumption?
19/11/2011 Q: How does one calculate amine loading for an amine absorber? (3)
18/11/2011 Q: Regarding the LPG Sulfrex Unit in RFCC, I have some questions.
We experience the increase of C4 sulfur content last Saturday (11/12) by the forming of the amine absorber(T-20701). ** a brief unit description is bottom of this writing: sulfur content of C4 goes up from 1~3 ppm to 16~18 ppm
Thus we replace the caustic of prewash drum(D-20702) & Extractor(T-20702). 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.
First is difference of amine inlet/outlet flow. Inlet lean amine flow is +6~8 m3/hr higher than outlet amine flow in amine absorber.
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.
Q1. Could you explain the reason for this phenomenon?
Q2. If amine quality is main cause, could you recommend the new guide of amine or other countermeasure?
**Brief LPG Sulfrex unit description :
LPG feed from R2R GAS Recovery unit is sent to the Amine absorber(T-20701). 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-20701 and rich amine is leaves the bottom of the absorber to amine regenerator.
LPG flows into the caustic prewash drum D-20702 for removal the last traces of H2S not removed in the amine absorber.
D-20703 is Caustic Settler. The settler drum allows to separate and return the entrained caustic to the oxidizer T-20703
09/11/2011 Q: I am have two diesel hydrotreaters in my refinery. One (35bar) is producing Euro III diesel with a sulfur specification of 350 ppm maximum and other (100bar) 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. (3)
28/10/2011 Q: How do we maintain the purity of Recycle hydrogen in a Diesel Hydrotreating unit? (5)
22/10/2011 Q: I am currently working in the Hydrotreater. 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? (5)
25/07/2011 Q: Lately we have been experienced frequent trip of Furnace in DHDS, we get positive draft and zero oxygen where this causes furnace to trip. Root cause? (3)
22/06/2011 Q: This is my NHT unit:
1 reactor with 1 fixed bed, volum 27m3, catalyst S-120 from UOP.
3 stage compressor: 4bar--> 10.5 bar--> 25 bar--> 43 bar.
Splitter 52 trays
Stripper 25 sieves trays.
If I change the feed for NHT unit from the SR naphtha with 100ppm wtS to the SR naphtha with 1230 ppm wt S ( because I changed the crude oil for DAV), what should I do to maintain the specification product for Platforming CCR (0.5 ppm S, 0.5 ppm N). And if I would like to revamp this unit for a product with 0.1 ppm wt S, what should I do?
20/06/2011 Q: We need to revamp our NHT. Before revamp: 23500 bpd SR Naphtha 100ppm S, Naphtha product for CCR feed has 0.5 wt ppm. After revamp: 30 000 bpd (90% vol SR Naphtha, 10 % coker Naphtha) with 0.1 wt ppm in product for our new regulation. We have 1 reactor (R1) with 1 bed of catalyst (18m3 catalyst in 27m3 reactor). I think we should install one more reactor. But I don't know which case is better between: Case 1: Feed-R1-R2-Stripper-splitter and Case 2: Feed-R1-Stripper-Splitter-R2 (recycle bottom product from splitter to R2)-R1. 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. It called SSRS Isocraking (single stage reverse sequencing), licensed by Chevron Lummus Global. In that article, they said that the revamped unit can run at 133% of original design capacity with the existing recycle gas compressor. 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. Unfortunately, It's not for NHT, It's Hydrocracking. Is it O.K if I use Case 2 for my NHT revamp?
18/05/2011 Q: Following is a brief overview of the problem we are currently facing at our Diesel Max Unit (Mild Hydrocracking Unit),
Our Diesel Max unit reactor having four beds, is equipped with three Quenches. MV for the 3rd Quench Gas Flow Valve started to increase and reached to a maximum value of 100% within 08 hrs. With increase in the MV opening of this Quench valve, flow across the valve remained consistent initially at around 4700 - 5000 Nm³/hr and then gradually started to decrease to a much lower value of 3,100 Nm³/hr at 100% opening at DCS at 70% Unit load.
Field observation was taken for the Quench valve and maximum opening found was 85% from field.
Field observation for any abnormal sound across the NRV was checked and found normal.
Similarly, Pressure drop across the reactor in the field on local PIs and across DPT,Delta T and Radial Spread across the reactor beds is observed and found no abnormality.
Actions Taken:
FT installed at the Quench valve was also drained and purged and found no error.
Unit load reduced to turn down ratio.
2nd Quench Gas flow at bed#3 increased to compensate for the reduction at Bed#4 Quench (3rd Quench).
Your expert opinion and guidance is requested on the Issue.
11/05/2011 Q: What is the standard value of sox/nox in atmosphere if emitting from hydrogen generation unit reformer for fg/naphtha/off gas firing? (1)
11/05/2011 Q: Our de-aerator conductivity is running high while de-aeration pressure is 0.3 kg/cm2g and temperature is 107 to 110 degree centigrade. Any thoughts on reasons and solutions? (3)
30/04/2011 Q: Recently a new DHDS Unit was successfully commissioned. At Unit Feed we have 25 micro meter Gas Oil SS cartridge filters. 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? What is the allowable limit of water in diesel Feed to DHDS Unit. Can high amount of residual water can choke filter elements? (2)
23/03/2011 Q: How do you calculate steam-to-carbon ratio in H.G.U.? (1)
19/03/2011 Q: Our NHT reactor DP increased when we changed the light feed material, orginal feed case design is 55% paraffins but we change this feed to 72% paraffinic feed. No reactor inlet temp / outlet temp or pressure changed. H2 consumption 100nm3/hr gone up but not too high. DP increased to 1500mmh2o from normal one (5000mmh2o).
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 1.1 to 1.2 only and no olefins with new feed case.
07/03/2011 Q: Recently we have 6 nos of recycle gas compressor trip incident, but even after 6 nos tripping we are not able to diagnose/analyse the reason for such trip. Tripping alarm in PLC for all occasion found to be same which is trip from compressor governor & trip from LCP (local control panel). Kindly suggest what may the probable reason for such a trip without any prior alarm? Compressor is steam driven.. RGC normal RPM ~11500 RPM. 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?
22/02/2011 Q: We have a Kerosene hydrotreater which is processing straight run light kerosene from crude unit to produce ATF. My feed conditions are : Temp at battery limit: 80-100 Deg C, pressure: 6.5 kg/cm2. Density: 0.804 @ 15 Deg C. Kerosene is being filtered by two basket type filter having 100 mesh (one stand by) (Filter temp around: 135-145 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.
what may be the reason of higher DP across filter?
Is there any chance of gum formation/ polymerisation (Because of additives in crude unit), which u can not see by naked eye, but may create DP?

Additional info:
Filter is getting chocked frequently. i.e. sometimes in 3-4 hrs (best achieved life 15-20 days). Once filter got chocked 16 times in 2.5 days. Dirty filter baskets are being cleaned by hydrojetting and followed by steaming.
Original design was of 25 micron (500 mesh), but because of frequent chocking filter mesh has been changed to 74 micron (200 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.
which feed characterization study can be carried out to identify problem.
16/01/2011 Q: Pyrolysis gasoline from Ethylene unit is sent to a recovery unit to recover C7 minus components. These are recovered in two columns under vacuum. Maximum temperature is at the bottom of the second column which is ~ 145 deg C. 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.
15/12/2010 Q: We have semi regeneration plateforming unit and currently it the third cycle. Catalyst is platinium/Rhenium.We are continuesly loosing Reactot Delta T. At SOR it was 129 deg C.while currently after 19 moths it is 90 deg C. MCH and CIS di methyl cyclo hexane in feed is 10.6 vol% while in reformate both are 1.2 vol%.Stab. ovhd gasses production is also very high. Hydrogen purity in previous cycles were above 90 while currently it is in between 87~88%. we have also experienced crude blend change in the current cycle. Naphthenes are currently 32 wt % while previously it was near to 35 wt %. such decrease in delta T was not experienced in previous two cycles. What are the possible causes?and what are the remedial actions to save the catalyst. (1)
24/11/2010 Q: We have SR Catalytic Reforming (Pt-Rh) Unit for 90.0 RONC production. It is our third cycle and the delta T of SR Reactors is decreasing rapidly day by day, but RONC is decreasing slightly or almost constant. 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.e; from 125 deg c to 88 deg C in 7-8 months), keeping in view that we are running plant at 110% load and our design H2 / HC ratio is 4.5 in first cycle. Is there any need of revision of H2/HC ration in third cycle, if yes then how?

Additional info: Its again me who put up the questions. 1-- Yes, it was text fault,, its Platinum - Rhenium. 2-- Please tell, what we have to check in feed and product regarding MCH? means which thing will proof us leakage in F/E? 3-- Their is only excessive increase in OVHD gases of stablizer. 4-- Hydrogen purity decreased from 90 to 85%. 5-- YES, H2 / HC ratio is easy to calculate, but i want to ask that during third cycle or as the cycles progress, is it necessary to revised this H2 /HC? if yes, then on what basis? 6-- We have increased RITs from 4-5 deg C but RON did not increase. 7-- We have decreased H2 / HC to about 3000 NM3/ hr and delta T improves from 89 deg C to 90 deg C. but a slight yellowish appearance of reformate was detected. ( what will be the reason?) But, RONC did not change
09/11/2010 Q: We have a hydrotreater, where I would like to limit my aromatics saturation. Can some one suggest if it is possible and what are the critical parameters for it? (5)
15/10/2010 Q: What would be a typical product distribution of Coker light gas oil (diesel range) on an FCC Unit when the Coker light gas oil has gone through a VGO Hydrotreater? (1)
13/10/2010 Q: Is it technically viable to inject sweet LPG ex CDU in reformate rundown (from bimetallic fixed bed reforming unit) with the sole purpose to increase reformate's RVP and RONC keeping in view that the reforming catalyst is nearing end of run and cannot meet the target sverity RONC? what are the repercussions? (6)
09/10/2010 Q: My company aims at further processing the atm. distillation residue (Mazot); and a hydrocracker unit has been chosen for this task. 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? (5)
07/10/2010 Q: Our Sour water stripper unit is a two stage operation. The first tower operates at 7 KSCg pressure and second tower operates at 0.8 KSCg pressure. 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, 100 C (Normal is 90C). Please suggest some solution. (2)
10/09/2010 Q: I am in Diesel Hydrotreater unit. Our feed contains maximum of 20% unsaturates (cycle oils). In the startup procedure, we were told that feed should be cut in at a reactor bed temperature of 260 C. But our current catalyst supplier has suggested that you can cut in feed even at 320 - 330 C. I just want to know what will be process implications if we cut in feed at 220C or 260C or 320C. Our diesel feed API is 33, feed sulfur is 1.2% wt, IBP 147 C and FBP 424 C.
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).
07/09/2010 Q: Our benzene product tank is internal floating roof tank with N2 blanketing which follow US EPA regulation. However measuring the VOC content at breath out shows as high at 15000 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?
01/09/2010 Q: I would like to know why there are two feed inlets in a Sour Water Stripping tower, but normally only one will be in use and the other will be blinded. In Sour water strippers why is chimney tray provided? (4)
31/08/2010 Q: I am working in a Diesel Hydrotreating unit. 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.
09/08/2010 Q: I work in Hydrocracking plant, where we commonly use turbine pump when running in normal condition, and backed up by motor pump as a spare pump.
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?
09/07/2010 Q: What is odorless kerosene? And how is it processed? (2)
07/07/2010 Q: Recently we carried out liquid phase sulfiding instead of gas phase sulfiding of freshly loaded hydrotreating & hydrocracker catalyst in the hydrocracker unit. Liquid phase sulfiding done with DMDS in light diesel oil & 50: 50 hydrogen/nitrogen pressure. After sulfiding phase over & unit feed cut-in with vacuum gas oil ex. vacuum distillation unit we encountered severe problem with sulfur in H2S form detected in light naphtha product (C5-135 deg C range) coming out from stabilizer column. Pl note fractionator column ovhd goes to D-ethanizer & stabilizer column after ovhd gas compression in three stg compressor to separate out fuel gas, LPG fraction & Light naphtha product.What may be the probable reason for H2S sulfur in high concentration (> 700 ppm) in light naphtha product? Is there any possibility of sulfur stripped out from liquid phase sufided catalyst? (2)
05/07/2010 Q: What will be the steam dew point at 93 degC and 1.1 kg/cm2g? 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 note total vapour mass flow in our distillation column ovhd is 274000 kg/hr and stripping steam flow (at 14 kg/cm2g & 220 degC) to column is 8490kg/hr. Overhead vapour is mixture of gaseous hydrocarbon (C1 to C5 range components) & stripping steam. Pl let me know if you need more data to answer my question. Our distillation column top operates at 1.1 kg/cm2g & 93 degC. (3)
03/07/2010 Q: What are the benefits of adding process steam in pre reformer inlet and reformer inlet separately? In some hydrogen plant it is mixed only in reformer inlet. What is the advantage of that? (2)
04/06/2010 Q: we offloaded our CCR catalyst for reactor checks/repairs. we want to reload and I don't know the maximum permissible coke allowed on the used catalyst. Should we load catalyst with coke level of 6%wt? (4)
26/05/2010 Q: In our Once Through Hydrocracker, the Fractionator Feed Furnace has options for both Fuel Oil and Fuel Gas Firing. 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? Another thing to consider, we are running at 70% T'Put and design conversion so in general the burners are supposed to operate at the given Heat Duty. (4)
18/05/2010 Q: I am currently using KBC Profimatics model to simulate hydrotreater reactors. Are there any other models available in the market? Are there any tools which can be helpful in daily monitoring of the hydrotreating reactors?  
13/05/2010 Q: Why is the range of electrical conductivity (unit:pico siemens/m) given as 50-450 for processing
jet-A1(kerosene fuels)?
what happen if range exceeds (i.e >520picosiemen/m) due to more chemical dosing during plant maloperations?
what does effects does it have on mechanical parts of aviation fuel?
13/05/2010 Q: What are the benefits of a top fired reformer versus a sided fired one? (5)
03/04/2010 Q: In Our Hydro Cracker we face the problem in Naphtha circuit. 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? (4)
19/03/2010 Q: I am working in DHDS unit. 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. The discharge of MGC (40 KSC) goes to suction of RGC (39.4 KSC). After starting the spare compressor and once it got 50% loaded, the make up gas rose from 25 Tons per day to 35 tons per days, simultaneously RGC amperage came down from 210 to 176 amps and discharge pressure of RGC came down from 61 KSC to 53 KSC and this dropping of amps and discharge pressure continued and unit tripped on low hydrogen pass flows. As the discharge pressure of RGC reduced the discharge flow also reduced. I didn't understand why the discharge pressure of RGC came down.

Additional Information: Separator pressure is constant and when the RGC tripped, it started raising. The suction flow was 265 Tons per day and when the RGC discharge pressure dropped, the suction flow also dropped to 255 Tons per day.
11/03/2010 Q: In a semi regen. fixed bed platforming unit for HOBC production, what is the impact of operating the unit at an increased Hydrogen to hydrocarbon mole ratio or increased H2 partial pressure than recommended? 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.
I want to know is there any positive or negative impact of this practice on reformate RONC, RVP and yield?
08/03/2010 Q: In our Once Through Hydrocracker Unit, the Recycle Gas Compressor is surging from 100% opening of the anti-surge valve to 0% without any change in process parameters. 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? (3)
22/02/2010 Q: In a two stage hydrocracker how can the selectivity for middle distillates can be improved at constant overall conversion? (2)
21/02/2010 Q: why is sulfiding of naphtha hydrotreater (NHT) catalyst necessary after carbon burning (regeneration) before normalization of NHT operation, when the main purpose of the catalyst itself is removal of sulfur from sour naphtha? (8)
19/02/2010 Q: I am working in DHDS. I would like to know the purpose of Carbon filter in Amine Recovery Unit. We use stripped water from Sour water stripping unit as wash water in DHDS over head coolers for dissolving ammonium salts. My query is if there are little amounts of ammonia and H2S in stripped water, and if we use the same stripped water in DHDS, will there be any problem in amine quality or will there be any effect in the quality of acid gas generated from ARU? We are facing the problem of increase in differential pressure across Carbon filter when we take stripped water in DHDS. (6)
14/02/2010 Q: We have a SR type reformer. The HDT catalyst (HR 306) was replaced by HR 506 after operating for 12 years and still getting DSN sulphur at about 0.1-0.2 ppm as reformer feed. After replacement of catalyst in April, 2009, the pressure drop was found to be increasing alarmingly and needed opening the reactors in Feb, 2010. On opening the reactor, a thick layer of about 1 ft of Fe dust was observed at the top of the reactor. We are planning to install a magnetic filter now to trap the Fe dust.
My query is with the same kind of feed and same type of loading (Sock) why the Fe dust deposition rate has increased to such an extent which was never experienced in the 12 years of operation from 1997 to 2009. The only process is that we are continuously dosing DMDS at the HDT to keep H2S at about 150-200ppm at HDT separator off gas as our Naphtha is very sweet in nature.
1. Can DMDS be the culprit for enhanced Fe deposit at the NHDT reactor?
2. In the latest catalyst loading, grading material was used which was not done in the previous loading. Has the new grading material caused the entire Fe-dust to be trapped at the reactor top.
3. What is the best location for installing a Magnetic filter?
4. Has the Fe-dust has increased due to the ageing Unit which was also observed during the last reformer section catalyst unloading in 2009.
07/02/2010 Q: For years, poor quality or hardprocess gasoline boiling range streams such as visbreaker has been a problem for refiners.
These materials contain such high quantities of di-olefins, in addition to sulfur and nitrogen compounds, that they are extremely difficult to process in conventional refinery units. The large di-olefin content of such streams renders them extremely reactive or unstable. If an attempt is made to simply hydrotreat these streams in a conventional hydrotreater, the reactive di-olefins form gum which plugs the conventional hydrotreating bed, or less frequently plugs the heat exchanger or heater upstream of the hydrotreating unit.
Based on above what are the options refiners have taken to solve this problem. In one refinery Naphtha is routed to Gas Concentration Unit before sending to Naphtha Hydrotreator Unit. GCU debutanizer reboiler tube bundle failed and heavy coke deposition observed.
Can Naphtha be routed directly to Naphtha Hydrotreator Unit ? If yes what will be demerits? What are the other options?
01/02/2010 Q: We are observing very high water content in the diesel storage tanks in our refinery. The rundown diesel from the diesel hydrotreater unit has a low water content of 0.05vol% which is within specification. Paraflow additive is injected downstream the DHT unit to improve the cold flow properties in winter and this is the only stream which mixes with the rundown diesel before being routed to tanks. Tanks are also observed to have longer than normal settling time to displace water.
Has any refiner experienced emulsification properties or increased water content in the diesel product due to addition of paraflow?
25/01/2010 Q: In a fixed bed semi regen. bimetallic Pt Re reformer catalyst with a target severity 90 RONC for HOBC production, as the cycle proceeds and catalyst ages, what is impact of catalyst ageing on GC of platformer stabilizer column (Debutanizer) off gases? (2)
31/12/2009 Q: In C5/C6 isomerization, we want to process low benzene (3%) and high C7+ (up to 15mol%). Our plant is designed to handle high benzene (up to 5.5%) and 7% of C7+ in the feed naphtha. We are expecting more severity on reactors and hence more cracking. If any one have experience on this, please share especially on yield and cracking. (what %age of isomerization yield will be lowered? and how much cracking will increase)  
10/12/2009 Q: What will be the effect of metal content (P, Na, Ca, Mg, Fe, Cu) in diesel feed on deactivation of DHDS or DHDT catalyst ? Whether demetallation catalyst used in hydrotreaters will be able to absorb these metals also?  
08/12/2009 Q: We have a catpoly hydrotreater that converts olefins to paraffins to produce petrol diesel and jet fuel. I just want to know the reaction/chemistry that should take place in the poly hydrotreater and the kinetics associated? (1)
15/11/2009 Q: What is PRD mode in automatic process control? (4)
01/11/2009 Q: What is main purpose of putting sealing steam in a turbine? (1)
01/11/2009 Q: In our DHDT recycle gas compressor primary seal vent flow at non driver end side has reduced to zero while it was previously 5 Nm3/hour. Driver end side flow is running between 30 Nm3/hour. What is the possible reason behind flow reduction? (1)
31/10/2009 Q: Refineries processing Crack Naphtha often face Naphtha Hydrotreating Reactor high pressure drop problems. At our refinery we do not have NHT Feed filtration. What are the major steps taken by refiners to prolong NHT operation? (7)
21/10/2009 Q: Is it safe to consider back pressure of 50-70 kg/cm2g when my PSV set pressure is at 229 kg/cm2g? Why are we limited to 3-5 kg/cm2g back pressure maximum when we are designing the HP flare? API 520 part 1 says that I can consider up to 50% of set pressure of balanced PSV, so can I consider up to 100 kg/cm2 g when my PSV is set at 220 kg/cm2g? If not, then what is the reason? (4)
18/09/2009 Q: In catalytic reforming unit what is the significance of N+2A, N+3.5A, and N+A in feed? What will be the minimum value to be required of N+2A for catalytic reforming unit either it is continuous catalytic reformer or semi regenerative reformer? (6)
18/09/2009 Q: We have naptha hydrotreater unit. At present unit is under construction phase and unit is being commissioned. We are going to be processing full range naphtha (C5 to 160 deg c) in our naphtha hydrotreater. What would be the recommended reactor inlet temperature and operating pressure for fresh feed cut in to reactor during normal start up? (5)
11/09/2009 Q: We have a butamer unit for isomerisation of nC4 into iC4. As our feed has over 2000ppm of olefins, there is an apprehension that u/s dryer beds will get fouled up and will call for early / premature replacements owing to potential polymerisation of olefins. Our questions are:-
1. Is there any refinery that processes with high olefins in the feed stock like ours, if yes, how they are managing the dryer life.
2. Did any refinery use hydroteatment for olefins saturation at such low level?
09/09/2009 Q: At the moment I am in the commissioning team of a Hydro cracking construction. We are still in the engineering phase of the project but soon we will start with on site field erection. As I am not familiar with Hydro crackers what is forum advice on inspection matters of this kind of units? What we should look closer in the disciplines of rotating, static and piping equipment. I know the main corrosion mechanisms in hydro crackers is ABS (ammonium bisulphide corrosion) and high temperature corrosion related with hydrogen but I would like to know from forum experience inspection point of view at what matters the commissioning inspectors must take more attention. (1)
18/08/2009 Q: Is there an an agreed percentage of sulphur that determines whether a crude is classed as low or high sulphur? (3)
15/08/2009 Q: In Our Recycle Gas compressor turbine seal steam pressure having too much fluctuation. Some time its pressure increase and some time decreases. What are the possible causes? (1)
12/08/2009 Q: In DHDT unit suppose benzene converted to cyclohexane and then cyclohexane converted to normal hexane. What is the mechanism of this reaction? How is aromatic converted to cyclohexane then how cyclohexane ring broken and converted to n-hexane? (3)
03/08/2009 Q: What is the feed-to-reflux ratio recommended in a naphtha hydrotreater stripper for efficient H2S stripping? (1)
21/07/2009 Q: For Euro-III diesel why must we maintain density 820 to 845 Kg/cube metre? How will performance be affected if this value is not maintained? (2)
21/07/2009 Q: Why must we maintain distillation of diesel 95% at 360 degrees centigrade for Euro-III ? If less or more what is the effect on engine performance? (2)
18/07/2009 Q: What is the basic difference between a thermal and pressure safety valve? (3)
17/07/2009 Q: Is straight run heavy naphtha treated in a merox unit fit to be used as feed in fixed bed bi metallic i.e. platinum rhenium based catalyst, platformer unit for octane improvement? (3)
02/07/2009 Q: What are the demerits of sending unstabilized reformate from fixed bed catalytic reformer unit directly to a storage tank? (7)
22/06/2009 Q: How can we establish the impact of high temperature water vapour on the compressor valve sealing element and its possible contribution to the melting of the element? (1)
05/06/2009 Q: What is the meaning of SOR (Start of Run ) and EOR (End of run) condition? Particularly in naphtha hydrotreating and c5/c6 isomerisation unit when we can say catalyst EOR condition is started? (1)
05/06/2009 Q: we are going to set up 1 MMTPA Naphtha hydrotreater in our refinery. Naphtha hydrotrater consist of vertical cylindrical Naphtha charge heater. The charge heater has two pass and each pass contain 32 nos of vertical tube. Our charge heater is in construction stage. For checking mechanical integrity of weld joints in tubes we are going to carry out hydrotest of tubes after placing tubes inside firebox. Also charge heater is part of reactor circuit. Any lumps of water will damage hydrotreater catalyst.
1) How should we remove water and debris from tubes after hydrotest as our tubes are vertical?
2) Are there any alternative methods to check weld integrity instead of hydrotest?
I know one option is pigging but we are not going to do it because of cost.
04/06/2009 Q: In what situation is a pneumatic test at one kg/cm2 to be preferred to a hydro test at the design pressure of a vessel? (2)
31/05/2009 Q: What are the different causes of pressure drop increment (gradually to maximum allowable limit) in a naphtha hydrotreater? (2)
28/03/2009 Q: What is lube oil supply temperature for any pump or compressor? Like feed pump, makeup gas and recycle gas compressor. (2)
25/03/2009 Q: Why do we need to maintain gas oil ratio in our diesel hydrotreater? (4)
19/03/2009 Q: With some experts projecting crude prices to creep back up to $75/bbl by mid-summer 2009, should we expect to see a higher level of refinery intermediates (e.g., heavy gas oil, "lifted" DAO, etc.) being exchanged among "networked" refining facilities?  
17/03/2009 Q: Are the declining costs of metallurgy providing an incentive for construction of 2000+ ton heavy-walled hydrocracking reactors? Is the application of advanced manufacturing techniques, such as Cr-Mo vanadium welding, becoming the 'norm' for fabrication of heavy walled hydrocracking reactors? What other developments coincide with new hydrocrackers designed to operate in a highly corrosive environment? (1)
12/03/2009 Q: We are experiencing excessive backwash frequency on our auto backwash filters (25 micron size) on the feed to the diesel hydrotreater unit.The hot feed is a blend of straight run kero, LDO and HDO which is fed directly from the crude unit with no makeup from intermediate storage. The feed when analyzed indicates a particulate level of about 6 ppm which in my opinion is low to cause such a problem. Has anyone experienced similar phenomena when the moisture levels in the feed are high? Moisture when analyzed was observed to be about 750 ppm in the feed. (3)
09/03/2009 Q: Why is a minimum circulation line not provided in some centrifugal pumps? For instance, in our stripper reflux pump it is provided, while in our diesel hydrotreater stripper it is not. (2)
04/03/2009 Q: What is the exact meaning high/low severity in case of refinery catalytic unit? (5)
02/03/2009 Q: In my DHDT (Diesel hydro treater unit), anti surge of recycle gas compressor remain open 20-25% always. Could anyone explain whether it is instrument fault or process problem? How can I rectify it? (3)
15/02/2009 Q: Why is the cetane index of diesel higher for high sulfur than low sulfur crude? (6)
15/02/2009 Q: What is the mechanism of aromatic saturation reaction in diesel hydrotreater reactor (i.e. step by step conversion from aromatic to paraffins)? (2)
07/02/2009 Q: What is the standard value of SOX & NOX in furnace stack outlet? Are the Values different in case of fuel oil firing and fuel gas firing? (3)
05/02/2009 Q: In a catalytic reforming unit the fines collection system may contain up to 30% catalyst pills, I would like to know what methods for fines/pills separation exists, along the lines of Density Grading to aid of optimum pills recovery. Also, is there is a better method?
What equipment is required and what are the physics involved?
03/02/2009 Q: Can anyone reference an article or research that comments on the effect lubricating oil from the makeup or recycle H2 compressors can have on catalyst life? (5)
01/02/2009 Q: What are the possible reasons for failure of silver corrosion test during ATF run in hydrotreater unit?
01/02/2009 Q: We have a plate heat exchanger as a reactor feed/effluent heat exchanger in DHT unit. this exchanger is very sensitive to debris/catalyst fine/ceramic ball chips.., accordingly a fine mesh (cone shape) is installed upstream the exchanger on the reactor effluent line. This filter is doing great by catching all scales and preventing them getting into the exchanger. However, when the Dp increases across the filter, we have to shutdown the unit and clean it. I'm looking for online cleaning, such as a dust collector, cyclone or whatever thing appropriate. the filter is installed in a piece that is same size as the piping and with no spare to avoid block valves in the reactor circuit. (3)
27/01/2009 Q: What are the chemical reactions taking place in diesel hydrotreater reactors that boost the cetane number, and how can these reactions be maintained? (3)
17/01/2009 Q: Can we reduce MPT time during startup of hdt unit? (In our case it is usually taken 28-30 hrs after M & I shutdown) (3)
17/01/2009 Q: We are facing a lot of back wash filter problems after 3-4 months in hydrotreater unit due to dirty or contaminated material carried along with feed so Is there any possibility of cleaning Backwash Filters during running without taking partial shutdown of unit? In other words I want to know is there any online method of cleaning filter? (8)
15/01/2009 Q: Which parameter - temperature or pressure - has more impact in a diesel hydrotreating unit in producing higher quality in terms of cetane no and product sulfur? (4)
16/12/2008 Q: What reliability issues can the use of high pressure unit charge pumps (multistage centrifugal pumps) in parallel pose to distillate hydrocracking processing ? (1)
09/12/2008 Q: antioxidant additives are used for aviation fuel with a maximum permitted dosage is 24 mg/l. What is the reason for this maximum value ?
what would be the consequences of adding more than 24 mg/l of antioxidant to jet fuel ?
16/10/2008 Q: Why is the non return valve fitted on the horizontal pipe line rather than the vertical one? (2)
23/09/2008 Q: I am working a project where I am trying detect phase changes. The project consist of detecting phase changes from water to butane by using flow meter density detectors. This idea is only for ideal case, but the reality is that, caustic may be present. Here is where the issue comes.
The question that I have is this: what method should I use to detect different phases. For example, mixed water and caustic? mixed Butane and Caustic? Again, the point is to detect phase density changes from water to butane.
08/09/2008 Q: Generally DMDS is used for hydrotreater catalyst sulfiding during the start up with new or regenerated catalyst. DBPS ( Di-t-butyle polysulfide) is known to be safer than DMDS due to its lower flash point than DMDS and other benefits as compared to DMDS. We like to know the prices of DBPS and DMDS. Which is costlier? (2)
06/08/2008 Q: Non Edible Vegetable oils contains metals like Ca, Mg, Si, Fe , P etc. and these vary from oil to oil and in ranges from 100 to 500 ppm.
We are looking for a process which can remove these metals to a level of <10 ppm. In addition the process should also work towards degumming of the oil.
04/08/2008 Q: How effective have membrane separation systems integrated into recent clean fuel strategies been in reducing sulphur levels, octane upgrading, etc.? (1)
02/08/2008 Q: What are the different types of high pressure exchangers and which ones among these are better than "Breech-lock" in terms of withstanding thermal shocks arising out of general power failure?

19/07/2008 Q: How one can derive Minimum Allowable Pressurization Temperature for a typical reactor? What is the procedure to work-out maximum allowable heat-up and cool-down rates for hydrocracking/treating reactor vessels? (2)
13/07/2008 Q: Can we process Heavy Aromatic Naphtha called "Reformate" Having aromatic concentration about 60 vol % into Hydrocracker unit to saturate aromatics ? Have anyone having experience of the same?  
11/07/2008 Q: How effective are the latest automation & control systems for ULSD hydrotreaters? Are they making a significant contribution in producing on-specification distillate product (< 8-10 ppm sulphur)? What is the feasibility of "extending" these control systems to upstream feed-stream distillation systems (i.e., tighter control of hard-to-remove refractory compounds entering hydrotreater)?  
10/07/2008 Q: are there any aromatic saturation happening in a Hydrocracker unit? What is the favourable condition for aromatic saturation? We have high aromatic naphtha (60 vol %). can we process it through hydrocracker for satration of aromatics? (1)
07/07/2008 Q: Under what circumstances is it cost effective to revamp the FCC main fractionator so that the amount of heavy FCC naphtha feed to ULSD hydrotreaters can be increased while still meeting finished ULSD product flash and distillation requirements? Are most ULSD hydrotreaters designed with a three-product stripper using a fired heater, or is a simple steam stripper adequate? (1)
01/07/2008 Q: We have sulphur guard bed in heavy naphtha stream which is going to catalytic reformer. The guard bed works on Chimisorbtion and is loaded with Nickel and aluminaosilicate based adsorbent. It reduces sulphur content to 0.1 wt ppm from incoming 0.5 wt ppm stream.
If Incoming naphtha stream sulphur content increased to 500 ppm, what is the expected life of sulphur guard bed? Also, to what extent can it remove sulphur from stream?
Are there special types of adsorbent available to cater for high amount of sulphur in incoming stream?
What is the average life of sulphur guard bed?
16/06/2008 Q: What is typical heat of reaction (in kcal/kmol of H2 consumed) for hydrodemetallization of vacuum gas oil? (1)
16/06/2008 Q: In a hydrotreater plant, water carry over in diesel is giving a problem. Is there any possibility that oxygenates are forming water in reaction? (3)
12/06/2008 Q: How are existing distillate hydrotreaters revamped to process higher volumes of feedstocks performing? What are some of the latest reactor and catalyst improvements that permit processing higher volumes of FCC LCO, coker naphtha or light coker gas oil through the distillate hydrotreater, and what are the corresponding benefits to downstream naphtha hydrotreater performance? (1)
19/05/2008 Q: We need some info about drying of hydrocracker catalyst by long period recycle gas circulation in case of start up and shut down of hydrocracker unit and problems caused by this phenomenon. Can anybody help us? Is it very harmful for catalysts? (1)
01/05/2008 Q: What are the conditions leading to brine production in a Catalyst cooler?  
26/04/2008 Q: What is the industry experience in handling coker gas oil in hydrocracker feed stock? How much absorption of HCGO ( as a wt% in total feed) has been achieved / being designed in existing units / revamps / new units ? Is there an optimum on HCGO portion in Hydrocracker feed mix ? Additional information on experience in critical equipment like Feed Filter / charge pump / Reactors/ RGC / Make Up Compressors / Fired Heaters/ Low pressure section / product treatment etc. would be appreciated. (1)
23/04/2008 Q: We have dual fired furnace. The FG generated in the process is used to run the furnace. Of course we switch over to fuel oil when such gas is not available, say, during start up. However of late, we have faced frequent burner blockage by carbon particles and sometimes a fireball coming out of the furnace. Due to this we are unable to run the furnace in fuel gas. However we have not noticed any carbon particle accumulation in the FG filters. Can anyone help us from similar kind of experience? (3)
18/04/2008 Q: what is selectivity and conversion in catalyst bed reactor and can anybody explain me about LHSV in reactor? (3)
08/04/2008 Q: How significant is the increase in hydrogen consumption in facilities where higher amounts of heavy VGO and heavy coker gas oil feeds are being treated in the gas oil hydrotreater? To what extent can catalyst selectivity help mitigate hydrogen consumption while treating these feeds? (1)
31/03/2008 Q: We have a conventional Thermal Hydrodealkylation plant (THDA) in our Petrochemicals complex, designed to primarily produce benzene for LAB production. The feed is taken from the refinery reformate stock. Our sister petrochemical company approached us to inquire about our capacity to produce cyclohexane fron our THDA.
We are writing to find out if such a proposal can be accommodated in the THDA unit and the required modifications to the existing facilities.A quick guide to anticipated changes to process operating parameters will be highly appreciated.
27/03/2008 Q: We have a corrosion problem in our hydrocracker unit high pressure fans (reactor effluent air coolers). There are three water pumps in the unit and by using one pump, water injection rate is 20m3/hr (by design). Recently, we encountered corrosion in the fan tubes and shut down unit five times in one year for repair. Sulfur and Nitrogen content of fresh feed is a little above design. Can anybody help us? Might it help if we increased water injection, using two pumps simultaneously? Has anybody experience in of this? (8)
24/03/2008 Q: We are experiencing falling of coke particles from the refinery hydrocarbon flare stack of late during sudden increase of gas flow subsequent to operation of dump valve of hydrocracker. We would like to know whether such incidents have occurred elsewhere ? If yes, what are the probable reasons and how can they be mitigated ? It may be noted that flare gas velocity during dump valve operation is well below 0.5 Mach. (1)
13/03/2008 Q: In our hydrocracker the reaction takes place only in the first bed. We're running on a reduced load. Is it possible to bring down the reaction to the next 3 beds and also to get a correct profile in the reactor? (1)
11/03/2008 Q: We don't have any clear procedure in our operating manual about degassing procedure for hydrocracker reactors. Can anyone help us?  
07/03/2008 Q: What is Best Practice for showing the Isolation
Valves for PGs, PTs,FTs, PDTs, LTs and Level Standpipes/Bridles in detailed engineering P&IDs?
04/03/2008 Q: We have a problem with our Hydrocracker VGO feed filters resulting in frequent backwash operations due to high Del P. Can you please ascertain the reason for the same as we do not get any FeS or suspended solids in the backwash stream analysis. Is it because of the asphaltenes as we process deep cut VGO (360-580+ degC) along with Heavy gas oil? (8)
03/03/2008 Q: Can we eliminate reactor outgasing step in a hydrocracker unit during shut down procedure? Is this action depend on reactor metallurgy? In new units, is this procedure necessary? (1)
28/02/2008 Q: While processing heavier and cracked feeds in Diesel Desulfurisation units the decativation could not take place due to metals poisoning or coke deposition. What are the views on predominant factor? If it is because of coke, is the only solution to make the feed lighter and process less of cracked stuff? However, if poisoning is due to metals, could a small bed of demet catalyst in the first bed prolong the life of the catalyst? (2)
20/02/2008 Q: In the query below, gas also contains CO2 which can help to maintain an acidic environment. This query is regarding an upstream processing facility. In a Sour Water stripper, maintaining pH of the water phase is essential for stripping. H2S tends to ionize in a basic environment, an acid environment is most conducive to keep H2S as an un-ionized form good for stripping. Published literature suggests maintaining a pH of 5 to 6 for H2S stripping.
Please note refinery sour gas generally also contains NH3 along with H2S and a desirable pH for stripping H2S and NH3 is 8 and is therefore different from the sour water above which does not come from a refinery and does not contain any NH3. I have the following questions in this regards,
a) For sour water containing only H2S and no NH3, what chemical is added prior to stripper for maintaining pH for good stripping? Also, is the chemical injection system similar to other chemical injection systems, eg Corrosion Inhibitor, i.e tank and pump?
b) Does partly ionized H2S not maintain its own pH without addition of chemical?
c) What is the typical column top operating pressure and is it maintained by a PCV? Are higher operating pressures any good for stripping considering the fact that they will reduce the column diameter?
17/02/2008 Q: What are the factors influencing the NHT catalyst performance towards nitrogen removal? And what is the most severe poison metal? (3)
07/02/2008 Q: Have hydraulic power recovery turbines (HPRTs) been included in any of the most recently completed or planned projects where core hydrocracking and amine regeneration is required? Besides energy savings incurred with the installation of an HPRT, how significant a role with HPRTs play in reducing CO2 emissions?  
01/02/2008 Q: How can I use prefractionation treatment to improve the quality of kerosene? (1)
11/01/2008 Q: What is temper embrittlement? What are the factors/parameters which affect it?  
08/01/2008 Q: Could the use of a 70 micron Element Filter (Wedge Wire), in place of a 25 micron Filter, for the filtration of hydrocracker unit feed, result in any problems? (4)
21/12/2007 Q: Is there any pilot plant scale-up data available WRT conversion of recalcitrant fibrous biomass materials into "reasonable" quality biocrudes? Do FCC or hydrotreating catalysts suppliers have any specific concerns WRT feeding small amounts of biocrudes into FCCUs or hydrotreaters? (1)
19/12/2007 Q: What is combined feed ratio? How is it calculated for a two stage hydrocracking process in a hydrocracker? (1)
27/11/2007 Q: How many, and what capacity, Gas To Liquid (GTL) plants are currently operational or under construction? (2)
09/11/2007 Q: We are looking for a chemical which can be used for removing Ni and V and also Fe from our hydrocracker unit feed (MVGO+LLC) by injection it to feed.
Also we think it is possible to eliminate these metals from crude oil source by adding chemicals to crude oil and removing metals in desalters.
Can anyone help us?
05/11/2007 Q: What are the practices followed at various refineries for hydroprocessing catalyst management. Is fresh catalyst charges or regeneration the preferred option? If regeneration is being followed, then for how many cycles? Are refineries maintaining stocks of different types of catalysts? (1)
01/11/2007 Q: please tell me some features of hydrocracking of heavy vacuum gas oil.  
01/11/2007 Q: What are the pre-requisites and requirements for a petrochemical plant start-up (e.g., naphtha-based steam cracker complex)? Does the facility’s effluent treatment plant need to be operational before actually feeding hydrocarbon into the complex? (2)
27/09/2007 Q: What are the practices followed for the disposal of Hydroprocessing catalysts:
1) Regeneration,
2) Metal recovery,
3) Disposal and replacement with new catalysts.
Is there any economic comparison of various options?
Who are the potential vendors working in different areas?
17/09/2007 Q: We have a conductivity loss problem in Jet Fuel produced from a Merox Unit.
Any experiences about the conductivity lost in Jet Fuel after Stadis 450 addition, will be very precious for the determining and finding the solution.
15/09/2007 Q: What processes are available for
1. the separation of oil from slack wax
2. the separation of wax fom residue wax
3. the hydrogenation of wax?
07/09/2007 Q: Outlet naphtha stream from tower in catalytic reforming unit that inlet in convection furnace can't shut down during catalyst regeneration. How we can solve this problem with least changing and regenerate catalyst and shut down tower at the same time? (4)
06/09/2007 Q: How I can calculate the production life of a polyethylene (gas phase) plant and, theoretically, what is the production life in years of Bp polyethylene gas phase technology polymerisation plant annual capacity 225,000 ton. HDPE and LLDPE?  
05/09/2007 Q: my question is related to an Isomerization unit. In a highly moisture sensitive Penex unit, what is the recommended practice to dry out the exchanger/system after the hydrotest of an exchanger in case of Leakage?
05/09/2007 Q: We were producing Aviation Turbine Fuel (ATF) by desulphurising 140-240 deg. cut kerosene through a Kerosene Hydro Desulphurisation Unit (KHDS). The electrical conductivity of ATF was maintained around 300 pS/m by injecting antistatic additive (Stadis 450) at a concentration of 1mg./lit in the product. Two years back we started producing ATF through the ATF Merox unit as well and are injecting Stadis 450 at the same rate from Merox Unit. But when the percentage of ATF produced from Merox Unit is going up to around 60% the conductivity of ATF drops down to around 150 pS/m (the marketing company require a minimum of 200 pS/m, even though the specification is 50-450 pS/m in ATF). What could the reasons for the drop in conductivity? (2)
04/09/2007 Q: What is the highest cut point (95% ASTM D1160) which can be done with Resid vacuum unit producing a high vacuum gasoil to feed a distillate single stage hydrocracking unit? What can be done in the hydrocracking unit to reduce the impact on the cycle length of such a heavy vacuum gasoil?  
04/09/2007 Q: Our hydrocracking unit is cycle length limited by the pressure drop build up in pretreating reactor. The pressure drop build up is caused by a deposition of iron sulfide at the top of the reactor. I want to know what can be done to solve this problem? (3)
17/08/2007 Q: In our naphtha hydrotreating unit, iron contents in stripper overhead boot are being reported on higher side for the last month. So far we have tried the following:
1. Increased corrosion inhibitor injection from 3 wt ppm (design) to 7 wt ppm.
2. Replaced the corrosion inhibitor
3. Cold condensate injection in reactor effluent increased from 3 to 5.5% of feed.
But iron contents are still high (2~3 ppm).
What could be the possible cause and what is the solution?
31/07/2007 Q: What are the processes available for removing 1,3 butadiene from Butene-1? Who are the licensors and what points should be considered in process selection? (1)
30/07/2007 Q: Fusel oil is generated in Crude Methanol purification/distillation. It has about 50% water and balance is various alcohols. To use it as a burner fuel is very taxing as the mass flow rate is very high compared to the fuel heat value.
The question is - How can this mixture be concentrated in an economic manner such that the water component is removed? This concentrated Fusel oil shall have much higher heat value and it shall improve burner firing efficiency.
28/07/2007 Q: How can process reconfigurations and reactor enhancements improve hydroprocessing catalyst performance? (3)
22/07/2007 Q: What potential opportunities are available for gasification of refinery residues? (1)